Oxidation System with Sidedraw Secondary Reactor

ABSTRACT

Disclosed is an optimized process and apparatus for more efficiently and economically carrying out the liquid-phase oxidation of an oxidizable compound. Such liquid-phase oxidation is carried out in a bubble column reactor that provides for a highly efficient reaction at relatively low temperatures. When the oxidized compound is para-xylene and the product from the oxidation reaction is crude terephthalic acid (CTA), such CTA product can be purified and separated by more economical techniques than could be employed if the CTA were formed by a conventional high-temperature oxidation process.

CROSS-REFERENCE TO RELATED APPLICATIONS

This Application is a continuation of application Ser. No. 11/846,846,entitled “Oxidation System with Sidedraw Secondary Reactor,” filed onAug. 29, 2007, which Application is a divisional of application Ser. No.11/366,005, filed on Mar. 1, 2006, the disclosure of which isincorporated herein by reference in its entirety.

FIELD OF THE INVENTION

This invention relates generally to a process for the production of apolycarboxylic acid composition. One aspect of the invention concernsthe partial oxidation of a dialkyl aromatic compound (e.g., para-xylene)to produce a crude aromatic dicarboxylic acid (e.g., crude terephthalicacid), which can thereafter be subjected to purification and separation.Another aspect of the invention concerns an improved reactor system thatprovides for a more effective and economical oxidation process.

BACKGROUND OF THE INVENTION

Oxidation reactions are employed in a variety of existing commercialprocesses. For example, liquid-phase oxidation is currently used for theoxidation of aldehydes to acids (e.g., propionaldehyde to propionicacid), the oxidation of cyclohexane to adipic acid, and the oxidation ofalkyl aromatics to alcohols, acids, or diacids. A particularlysignificant commercial oxidation process in the latter category(oxidation of alkyl aromatics) is the liquid-phase catalytic partialoxidation of para-xylene to terephthalic acid. Terephthalic acid is animportant compound with a variety of applications. The primary use ofterephthalic acid is as a feedstock in the production of polyethyleneterephthalate (PET). PET is a well-known plastic used in greatquantities around the world to make products such as bottles, fibers,and packaging.

In a typical liquid-phase oxidation process, including partial oxidationof para-xylene to terephthalic acid, a liquid-phase feed stream and agas-phase oxidant stream are introduced into a reactor and form amulti-phase reaction medium in the reactor. The liquid-phase feed streamintroduced into the reactor contains at least one oxidizable organiccompound (e.g., para-xylene), while the gas-phase oxidant streamcontains molecular oxygen. At least a portion of the molecular oxygenintroduced into the reactor as a gas dissolves into the liquid phase ofthe reaction medium to provide oxygen availability for the liquid-phasereaction. If the liquid phase of the multi-phase reaction mediumcontains an insufficient concentration of molecular oxygen (i.e., ifcertain portions of the reaction medium are “oxygen-starved”),undesirable side-reactions can generate impurities and/or the intendedreactions can be retarded in rate. If the liquid phase of the reactionmedium contains too little of the oxidizable compound, the rate ofreaction may be undesirably slow. Further, if the liquid phase of thereaction medium contains an excess concentration of the oxidizablecompound, additional undesirable side-reactions can generate impurities.

Conventional liquid-phase oxidation reactors are equipped with agitationmeans for mixing the multi-phase reaction medium contained therein.Agitation of the reaction medium is supplied in an effort to promotedissolution of molecular oxygen into the liquid phase of the reactionmedium, maintain relatively uniform concentrations of dissolved oxygenin the liquid phase of the reaction medium, and maintain relativelyuniform concentrations of the oxidizable organic compound in the liquidphase of the reaction medium.

Agitation of the reaction medium undergoing liquid-phase oxidation isfrequently provided by mechanical agitation means in vessels such as,for example, continuous stirred tank reactors (CSTRs). Although CSTRscan provide thorough mixing of the reaction medium, CSTRs have a numberof drawbacks. For example, CSTRs have a relatively high capital cost dueto their requirement for expensive motors, fluid-sealed bearings anddrive shafts, and/or complex stirring mechanisms. Further, the rotatingand/or oscillating mechanical components of conventional CSTRs requireregular maintenance. The labor and shutdown time associated with suchmaintenance adds to the operating cost of CSTRs. However, even withregular maintenance, the mechanical agitation systems employed in CSTRsare prone to mechanical failure and may require replacement overrelatively short periods of time.

Bubble column reactors provide an attractive alternative to CSTRs andother mechanically agitated oxidation reactors. Bubble column reactorsprovide agitation of the reaction medium without requiring expensive andunreliable mechanical equipment. Bubble column reactors typicallyinclude an elongated upright reaction zone within which the reactionmedium is contained. Agitation of the reaction medium in the reactionzone is provided primarily by the natural buoyancy of gas bubbles risingthrough the liquid phase of the reaction medium. This natural-buoyancyagitation provided in bubble column reactors reduces capital andmaintenance costs relative to mechanically agitated reactors. Further,the substantial absence of moving mechanical parts associated withbubble column reactors provides an oxidation system that is less proneto mechanical failure than mechanically agitated reactors.

When liquid-phase partial oxidation of para-xylene is carried out in aconventional oxidation reactor (CSTR or bubble column), the productwithdrawn from the reactor is typically a slurry comprising crudeterephthalic acid (CTA) and a mother liquor. CTA contains relativelyhigh levels of impurities (e.g., 4-carboxybenzaldehyde, para-toluicacid, fluorenones, and other color bodies) that render it unsuitable asa feedstock for the production of PET. Thus, the CTA produced inconventional oxidation reactors is typically subjected to a purificationprocess that converts the CTA into purified terephthalic acid (PTA)suitable for making PET.

One typical purification process for converting CTA to PTA includes thefollowing steps: (1) replacing the mother liquor of the CTA-containingslurry with water, (2) heating the CTA/water slurry to dissolve the CTAin water, (3) catalytically hydrogenating the CTA/water solution toconvert impurities to more desirable and/or easily-separable compounds,(4) precipitating the resulting PTA from the hydrogenated solution viamultiple crystallization steps, and (5) separating the crystallized PTAfrom the remaining liquids. Although effective, this type ofconventional purification process can be very expensive. Individualfactors contributing to the high cost of conventional CTA purificationmethods include, for example, the heat energy required to promotedissolution of the CTA in water, the catalyst required forhydrogenation, the hydrogen stream required for hydrogenation, the yieldloss caused by hydrogenation of some terephthalic acid, and the multiplevessels required for multi-step crystallization. Thus, it would bedesirable to provide an oxidation system capable of producing a CTAproduct that could be purified without requiring heat-promoteddissolution in water, hydrogenation, and/or multi-step crystallization.

OBJECTS OF THE INVENTION

It is, therefore, an object of the present invention to provide a moreeffective and economical liquid-phase oxidation system.

Another object of the invention is to provide a more effective andeconomical reactor and process for the liquid-phase catalytic partialoxidation of para-xylene to terephthalic acid.

Still another object of the invention is to provide a bubble columnreactor that facilitates improved liquid-phase oxidation reactions withreduced formation of impurities.

Yet another object of the invention is to provide a more effective andeconomical system for producing pure terephthalic acid (PTA) vialiquid-phase oxidation of para-xylene to produce crude terephthalic acid(CTA) and subsequently, purifying the CTA to PTA.

A further object of the invention is to provide a bubble column reactorfor oxidizing para-xylene and producing a CTA product capable of beingpurified without requiring heat-promoted dissolution of the CTA inwater, hydrogenation of the dissolved CTA, and/or multi-stepcrystallization of the hydrogenated PTA.

It should be noted that the scope of the present invention, as definedin the appended claims, is not limited to processes or apparatusescapable of realizing all of the objects listed above. Rather, the scopeof the claimed invention may encompass a variety of systems that do notaccomplish all or any of the above-listed objects. Additional objectsand advantages of the present invention will be readily apparent to oneskilled in the art upon reviewing the following detailed description andassociated drawings.

SUMMARY OF THE INVENTION

One embodiment of the present invention concerns a process for making apolycarboxylic acid composition, the process comprising the followingsteps: (a) subjecting a multi-phase reaction medium to oxidation in aprimary oxidation reactor to thereby produce a first slurry; and (b)subjecting at least a portion of the first slurry to further oxidationin a secondary oxidation reactor, wherein the secondary oxidationreactor is a bubble column reactor.

Another embodiment of the present invention concerns a reactor system.The reactor system includes a primary oxidation reactor and a secondaryoxidation reactor. The primary oxidation reactor defines a first inletand a first outlet. The secondary oxidation reactor is a bubble columnreactor that defines a second inlet and a second outlet. The firstoutlet is coupled in fluid flow communication with the second inlet.

BRIEF DESCRIPTION OF THE DRAWINGS

Preferred embodiments of the invention are described in detail belowwith reference to the attached drawing figures, wherein;

FIG. 1 is a side view of an oxidation reactor constructed in accordancewith one embodiment of the present invention, particularly illustratingthe introduction of feed, oxidant, and reflux streams into the reactor,the presence of a multi-phase reaction medium in the reactor, and thewithdrawal of a gas and a slurry from the top and bottom of the reactor,respectively;

FIG. 2 is an enlarged sectional side view of the bottom of the bubblecolumn reactor taken along line 2-2 in FIG. 3, particularly illustratingthe location and configuration of a oxidant sparger used to introducethe oxidant stream into the reactor;

FIG. 3 is a top view of the oxidant sparger of FIG. 2, particularlyillustrating that there are no oxidant discharge openings in the top ofthe oxidant sparger;

FIG. 4 is a bottom view of the oxidant sparger of FIG. 2, particularlyillustrating the configuration of the oxidant discharge openings in thebottom of the oxidant sparger;

FIG. 5 is a sectional side view of the oxidant sparger taken along line5-5 in FIG. 3, particularly illustrating the orientation of the oxidantdischarge openings in the bottom of the oxidant sparger; FIG. 6 is anenlarged side view of the bottom portion of the bubble column reactor,particular illustrating a system for introducing the feed stream intothe reactor at multiple, vertically-space locations;

FIG. 7 is a sectional top view taken along line 7-7 in FIG. 6,particularly illustrating how the feed introduction system shown in FIG.6 distributes the feed stream into in a preferred radial feed zone (FZ)and more than one azimuthal quadrant (Q₁, Q₂, Q₃, Q₄);

FIG. 8 is a sectional top view similar to FIG. 7, but illustrating analternative means for discharging the feed stream into the reactor usingbayonet tubes each having a plurality of small feed openings;

FIG. 9 is an isometric view of an alternative system for introducing thefeed stream into the reaction zone at multiple vertically-spacelocations without requiring multiple vessel penetrations, particularlyillustrating that the feed distribution system can be at least partlysupported on the oxidant sparger;

FIG. 10 is a side view of the single-penetration feed distributionsystem and oxidant sparger illustrated in FIG. 9;

FIG. 11 is a sectional top view taken along line 11-11 in FIG. 10 andfurther illustrating the single-penetration feed distribution systemsupported on the oxidant sparger;

FIG. 12 is a side view of a bubble column reactor equipped with internaland external reaction vessels;

FIG. 13 is an enlarged sectional view of the bubble column reactor ofFIG. 12 taken along line 13-13, particularly illustrating the relativeorientation of the internal and external reaction vessels;

FIG. 14 is a side view of an alternative bubble column reactor equippedwith internal and external reaction vessels, particularly illustratingthat the external reaction vessel has a stepped diameter;

FIG. 15 is a side view of a bubble column reactor equipped with anexternal secondary oxidation reactor that receives a slurry from asidedraw in the primary oxidation reactor;

FIG. 16 is a side view of a bubble column reactor equipped with anopen-ended external secondary oxidation reactor that receives slurryfrom an enlarged opening in the side of the primary oxidation reactor;

FIG. 17 a is a schematic side view of a bubble column reactor equippedwith an internal structure for enhancing the hydrodynamics of thereactor;

FIG. 17 b is a sectional view of the reactor of FIG. 17 a taken alongline 17 b-17 b in FIG. 17 a;

FIG. 18 a is a schematic side view of a bubble column reactor equippedwith a first alternative internal structure for enhancing thehydrodynamics of the reactor;

FIG. 18 b is a sectional view of the reactor of FIG. 18 a taken alongline 18 b-18 b in FIG. 18 a;

FIG. 19 a is a schematic side view of a bubble column reactor equippedwith a second alternative internal structure for enhancing thehydrodynamics of the reactor;

FIG. 19 b is a sectional view of the reactor of FIG. 19 a taken alongline 19 b-19 b in FIG. 19 a;

FIG. 20 a is a schematic side view of a bubble column reactor equippedwith a third alternative internal structure for enhancing thehydrodynamics of the reactor;

FIG. 20 b is a sectional view of the reactor of FIG. 20 a taken alongline 20 b-20 b in FIG. 20 a;

FIG. 21 a is a schematic side view of a bubble column reactor equippedwith a fourth alternative internal structure for enhancing thehydrodynamics of the reactor;

FIG. 21 b is a sectional view of the reactor of FIG. 21 a taken alongline 21 b-21 b in FIG. 21 a;

FIG. 22 a is a schematic side view of a bubble column reactor equippedwith a fifth alternative internal structure for enhancing thehydrodynamics of the reactor;

FIG. 22 b is a sectional view of the reactor of FIG. 22 a taken alongline 22 b-22 b in FIG. 22 a;

FIG. 23 a is a schematic side view of a bubble column reactor equippedwith a sixth alternative internal structure for enhancing thehydrodynamics of the reactor;

FIG. 23 b is a sectional view of the reactor of FIG. 23 a taken alongline 23 b-23 b in FIG. 23 a;

FIG. 24 a is a schematic side view of a bubble column reactor equippedwith a seventh alternative internal structure for enhancing thehydrodynamics of the reactor;

FIG. 24 b is a sectional view of the reactor of FIG. 24 a taken alongline 24 b-24 b in FIG. 24 a;

FIG. 25 a is a schematic view of a stepped-diameter bubble columnreactor equipped with a hydrodynamic-enhancing internal structure;

FIG. 25 b is a sectional view of the reactor of FIG. 25 a taken alongline 25 b-25 b in FIG. 25 a;

FIG. 26 is a side view of a bubble column reactor containing amulti-phase reaction medium, particularly illustrating the reactionmedium being theoretically partitioned into 30 horizontal slices ofequal volume in order to quantify certain gradients in the reactionmedium;

FIG. 27 is a side view of a bubble column reactor containing amulti-phase reaction medium, particularly illustrating first and seconddiscrete 20-percent continuous volumes of the reaction medium that havesubstantially different oxygen concentrations and/or oxygen consumptionrates;

FIGS. 28A and 28B are magnified views of crude terephthalic acid (CTA)particles produced in accordance with one embodiment of the presentinvention, particularly illustrating that each CTA particle is a lowdensity, high surface area particle composed of a plurality ofloosely-bound CTA sub-particles;

FIGS. 29A and 29B are magnified views of a conventionally-produced CTA,particularly illustrating that the conventional CTA particle has alarger particle size, higher density, and lower surface area than theinventive CTA particle of FIGS. 28A and 28B;

FIG. 30 is a simplified process flow diagram of a prior art process formaking purified terephthalic acid (PTA); and

FIG. 31 is a simplified process flow diagram of a process for making PTAin accordance with one embodiment of the present invention.

DETAILED DESCRIPTION

One embodiment of the present invention concerns the liquid-phasepartial oxidation of an oxidizable compound. Such oxidation ispreferably carried out in the liquid phase of a multi-phase reactionmedium contained in one or more agitated reactors. Suitable agitatedreactors include, for example, bubble-agitated reactors (e.g., bubblecolumn reactors), mechanically agitated reactors (e.g., continuousstirred tank reactors), and flow agitated reactors (e.g., jet reactors).In one embodiment of the invention, the liquid-phase oxidation iscarried out using at least one bubble column reactor.

As used herein, the term “bubble column reactor” shall denote a reactorfor facilitating chemical reactions in a multi-phase reaction medium,wherein agitation of the reaction medium is provided primarily by theupward movement of gas bubbles through the reaction medium. As usedherein, the term “agitation” shall denote work dissipated into thereaction medium causing fluid flow and/or mixing. As used herein, theterms “majority,” “primarily,” and “predominately” shall mean more than50 percent. As used herein, the term “mechanical agitation” shall denoteagitation of the reaction medium caused by physical movement of a rigidor flexible element(s) against or within the reaction medium. Forexample, mechanical agitation can be provided by rotation, oscillation,and/or vibration of internal stirrers, paddles, vibrators, or acousticaldiaphragms located in the reaction medium. As used herein, the term“flow agitation” shall denote agitation of the reaction medium caused byhigh velocity injection and/or recirculation of one or more fluids inthe reaction medium. For example, flow agitation can be provided bynozzles, ejectors, and/or eductors.

In a preferred embodiment of the present invention, less than about 40percent of the agitation of the reaction medium in the bubble columnreactor during oxidation is provided by mechanical and/or flowagitation, more preferably less than about 20 percent of the agitationis provided by mechanical and/or flow agitation, and most preferablyless than 5 percent of the agitation is provided by mechanical and/orflow agitation. Preferably, the amount of mechanical and/or flowagitation imparted to the multi-phase reaction medium during oxidationis less than about 3 kilowatts per cubic meter of the reaction medium,more preferably less than about 2 kilowatts per cubic meter, and mostpreferably less than 1 kilowatt per cubic meter.

Referring now to FIG. 1, a preferred bubble column reactor 20 isillustrated as comprising a vessel shell 22 having a reaction section 24and a disengagement section 26. Reaction section 24 defines a reactionzone 28, while disengagement section 26 defines a disengagement zone 30.A predominately liquid-phase feed stream is introduced into reactionzone 28 via feed inlets 32 a,b,c,d. A predominately gas-phase oxidantstream is introduced into reaction zone 28 via an oxidant sparger 34located in the lower portion of reaction zone 28. The liquid-phase feedstream and gas-phase oxidant stream cooperatively form a multi-phasereaction medium 36 within reaction zone 28. Multi-phase reaction medium36 comprises a liquid phase and a gas phase. More preferably, multiphasereaction medium 36 comprises a three-phase medium having solid-phase,liquid-phase, and gas-phase components. The solid-phase component of thereaction medium 36 preferably precipitates within reaction zone 28 as aresult of the oxidation reaction carried out in the liquid phase ofreaction medium 36. Bubble column reactor 20 includes a slurry outlet 38located near the bottom of reaction zone 28 and a gas outlet 40 locatednear the top of disengagement zone 30. A slurry effluent comprisingliquid-phase and solid-phase components of reaction medium 36 iswithdrawn from reaction zone 28 via slurry outlet 38, while apredominantly gaseous effluent is withdrawn from disengagement zone 30via gas outlet 40. The liquid-phase feed stream introduced into bubblecolumn reactor 20 via feed inlets 32 a,b,c,d preferably comprises anoxidizable compound, a solvent, and a catalyst system.

The oxidizable compound present in the liquid-phase feed streampreferably comprises at least one hydrocarbyl group. More preferably,the oxidizable compound is an aromatic compound. Still more preferably,the oxidizable compound is an aromatic compound with at least oneattached hydrocarbyl group or at least one attached substitutedhydrocarbyl group or at least one attached heteroatom or at least oneattached carboxylic acid function (—COOH). Even more preferably, theoxidizable compound is an aromatic compound with at least one attachedhydrocarbyl group or at least one attached substituted hydrocarbyl groupwith each attached group comprising from 1 to 5 carbon atoms. Yet stillmore preferably, the oxidizable compound is an aromatic compound havingexactly two attached groups with each attached group comprising exactlyone carbon atom and consisting of methyl groups and/or substitutedmethyl groups and/or at most one carboxylic acid group. Even still morepreferably, the oxidizable compound is para-xylene, meta-xylene,para-tolualdehyde, meta-tolualdehyde, para-toluic acid, meta-toluicacid, and/or acetaldehyde. Most preferably, the oxidizable compound ispara-xylene.

A “hydrocarbyl group,” as defined herein, is at least one carbon atomthat is bonded only to hydrogen atoms or to other carbon atoms. A“substituted hydrocarbyl group,” as defined herein, is at least onecarbon atom bonded to at least one heteroatom and to at least onehydrogen atom. “Heteroatoms,” as defined herein, are all atoms otherthan carbon and hydrogen atoms. Aromatic compounds, as defined herein,comprise an aromatic ring, preferably having at least 6 carbon atoms,even more preferably having only carbon atoms as part of the ring.Suitable examples of such aromatic rings include, but are not limitedto, benzene, biphenyl, terphenyl, naphthalene, and other carbon-basedfused aromatic rings.

If the oxidizable compound present in the liquid-phase feed stream is anormally-solid compound (i.e., is a solid at standard temperature andpressure), it is preferred for the oxidizable compound to besubstantially dissolved in the solvent when introduced into reactionzone 28. It is preferred for the boiling point of the oxidizablecompound at atmospheric pressure to be at least about 50° C. Morepreferably, the boiling point of the oxidizable compound is in the rangeof from about 80 to about 400° C., and most preferably in the range offrom 125 to 155° C. The amount of oxidizable compound present in theliquid-phase feed is preferably in the range of from about 2 to about 40weight percent, more preferably in the range of from about 4 to about 20weight percent, and most preferably in the range of from 6 to 15 weightpercent.

It is now noted that the oxidizable compound present in the liquid-phasefeed may comprise a combination of two or more different oxidizablechemicals. These two or more different chemical materials can be fedcommingled in the liquid-phase feed stream or may be fed separately inmultiple feed streams. For example, an oxidizable compound comprisingpara-xylene, meta-xylene, para-tolualdehyde, para-toluic acid, andacetaldehyde may be fed to the reactor via a single inlet or multipleseparate inlets.

The solvent present in the liquid-phase feed stream preferably comprisesan acid component and a water component. The solvent is preferablypresent in the liquid-phase feed stream at a concentration in the rangeof from about 60 to about 98 weight percent, more preferably in therange of from about 80 to about 96 weight percent, and most preferablyin the range of from 85 to 94 weight percent. The acid component of thesolvent is preferably primarily an organic low molecular weightmonocarboxylic acid having 1-6 carbon atoms, more preferably 2 carbonatoms. Most preferably, the acid component of the solvent is primarilyacetic acid. Preferably, the acid component makes up at least about 75weight percent of the solvent, more preferably at least about 80 weightpercent of the solvent, and most preferably 85 to 98 weight percent ofthe solvent, with the balance being primarily water. The solventintroduced into bubble column reactor 20 can include small quantities ofimpurities such as, for example, para-tolualdehyde, terephthaldehyde,4-carboxybenzaldehyde (4-CBA), benzoic acid, para-toluic acid,para-toluic aldehyde, alpha-bromo-para-toluic acid, isophthalic acid,phthalic acid, trimellitic acid, polyaromatics, and/or suspendedparticulate. It is preferred that the total amount of impurities in thesolvent introduced into bubble column reactor 20 is less than about 3weight percent.

The catalyst system present in the liquid-phase feed stream ispreferably a homogeneous, liquid-phase catalyst system capable ofpromoting oxidation (including partial oxidation) of the oxidizablecompound. More preferably, the catalyst system comprises at least onemultivalent transition metal. Still more preferably, the multivalenttransition metal comprises cobalt. Even more preferably, the catalystsystem comprises cobalt and bromine Most preferably, the catalyst systemcomprises cobalt, bromine, and manganese.

When cobalt is present in the catalyst system, it is preferred for theamount of cobalt present in the liquid-phase feed stream to be such thatthe concentration of cobalt in the liquid phase of reaction medium 36 ismaintained in the range of from about 300 to about 6,000 parts permillion by weight (ppmw), more preferably in the range of from about 700to about 4,200 ppmw, and most preferably in the range of from 1,200 to3,000 ppmw. When bromine is present in the catalyst system, it ispreferred for the amount of bromine present in the liquid-phase feedstream to be such that the concentration of bromine in the liquid phaseof reaction medium 36 is maintained in the range of from about 300 toabout 5,000 ppmw, more preferably in the range of from about 600 toabout 4,000 ppmw, and most preferably in the range of from 900 to 3,000ppmw. When manganese is present in the catalyst system, it is preferredfor the amount of manganese present in the liquid-phase feed stream tobe such that the concentration of manganese in the liquid phase ofreaction medium 36 is maintained in the range of from about 20 to about1,000 ppmw, more preferably in the range of from about 40 to about 500ppmw, most preferably in the range of from 50 to 200 ppmw.

The concentrations of the cobalt, bromine, and/or manganese in theliquid phase of reaction medium 36, provided above, are expressed on atime-averaged and volume-averaged basis. As used herein, the term“time-averaged” shall denote an average of at least 10 measurementstaken equally over a continuous period of at least 100 seconds. As usedherein, the term “volume-averaged” shall denote an average of at least10 measurements taken at uniform 3-dimensional spacing throughout acertain volume.

The weight ratio of cobalt to bromine (Co:Br) in the catalyst systemintroduced into reaction zone 28 is preferably in the range of fromabout 0.25:1 to about 4:1, more preferably in the range of from about0.5:1 to about 3:1, and most preferably in the range of from 0.75:1 to2:1. The weight ratio of cobalt to manganese (Co:Mn) in the catalystsystem introduced into reaction zone 28 is preferably in the range offrom about 0.3:1 to about 40:1, more preferably in the range of fromabout 5:1 to about 30:1, and most preferably in the range of from 10:1to 25:1.

The liquid-phase feed stream introduced into bubble column reactor 20can include small quantities of impurities such as, for example,toluene, ethylbenzene, para-tolualdehyde, terephthaldehyde,4-carboxybenzaldehyde (4-CBA), benzoic acid, para-toluic acid,para-toluic aldehyde, alpha bromo para-toluic acid, isophthalic acid,phthalic acid, trimellitic acid, polyaromatics, and/or suspendedparticulate. When bubble column reactor 20 is employed for theproduction of terephthalic acid, meta-xylene and ortho-xylene are alsoconsidered impurities. It is preferred that the total amount ofimpurities in the liquid-phase feed stream introduced into bubble columnreactor 20 is less than about 3 weight percent.

Although FIG. 1 illustrates an embodiment where the oxidizable compound,the solvent, and the catalyst system are mixed together and introducedinto bubble column reactor 20 as a single feed stream, in an alternativeembodiment of the present invention, the oxidizable compound, thesolvent, and the catalyst can be separately introduced into bubblecolumn reactor 20. For example, it is possible to feed a purepara-xylene stream into bubble column reactor 20 via an inlet separatefrom the solvent and catalyst inlet(s).

The predominately gas-phase oxidant stream introduced into bubble columnreactor 20 via oxidant sparger 34 comprises molecular oxygen (O₂).Preferably, the oxidant stream comprises in the range of from about 5 toabout 40 mole percent molecular oxygen, more preferably in the range offrom about 15 to about 30 mole percent molecular oxygen, and mostpreferably in the range of from 18 to 24 mole percent molecular oxygen.It is preferred for the balance of the oxidant stream to be comprisedprimarily of a gas or gasses, such as nitrogen, that are inert tooxidation. More preferably, the oxidant stream consists essentially ofmolecular oxygen and nitrogen. Most preferably, the oxidant stream isdry air that comprises about 21 mole percent molecular oxygen and about78 to about 81 mole percent nitrogen. In an alternative embodiment ofthe present invention, the oxidant stream can comprise substantiallypure oxygen.

Referring again to FIG. 1, bubble column reactor 20 is preferablyequipped with a reflux distributor 42 positioned above an upper surface44 of reaction medium 36. Reflux distributor 42 is operable to introducedroplets of a predominately liquid-phase reflux stream intodisengagement zone 30 by any means of droplet formation known in theart. More preferably, reflux distributor 42 produces a spray of dropletsdirected downwardly towards upper surface 44 of reaction medium 36.Preferably, this downward spray of droplets affects (i.e., engages andinfluences) at least about 50 percent of the maximum horizontalcross-sectional area of disengagement zone 30. More preferably, thespray of droplets affects at least about 75 percent of the maximumhorizontal cross-sectional area of disengagement zone 30. Mostpreferably, the spray of droplets affects at least 90 percent of themaximum horizontal cross-sectional area of disengagement zone 30. Thisdownward liquid reflux spray can help prevent foaming at or above uppersurface 44 of reaction medium 36 and can also aid in the disengagementof any liquid or slurry droplets entrained in the upwardly moving gasthat flows towards gas outlet 40. Further, the liquid reflux may serveto reduce the amount of particulates and potentially precipitatingcompounds (e.g., dissolved benzoic acid, para-toluic acid, 4-CBA,terephthalic acid, and catalyst metal salts) exiting in the gaseouseffluent withdrawn from disengagement zone 30 via gas outlet 40. Inaddition, the introduction of reflux droplets into disengagement zone 30can, by a distillation action, be used to adjust the composition of thegaseous effluent withdrawn via gas outlet 40.

The liquid reflux stream introduced into bubble column reactor 20 viareflux distributor 42 preferably has about the same composition as thesolvent component of the liquid-phase feed stream introduced into bubblecolumn reactor 20 via feed inlets 32 a,b,c,d. Thus, it is preferred forthe liquid reflux stream to comprise an acid component and water. Theacid component of the reflux stream is preferably a low molecular weightorganic monocarboxylic acid having 1-6 carbon atoms, more preferably 2carbon atoms. Most preferably, the acid component of the reflux streamis acetic acid. Preferably, the acid component makes up at least about75 weight percent of the reflux stream, more preferably at least about80 weight percent of the reflux stream, and most preferably 85 to 98weight percent of the reflux stream, with the balance being water.Because the reflux stream typically has substantially the samecomposition as the solvent in the liquid-phase feed stream, when thisdescription refers to the “total solvent” introduced into the reactor,such “total solvent” shall include both the reflux stream and thesolvent portion of the feed stream.

During liquid-phase oxidation in bubble column reactor 20, it ispreferred for the feed, oxidant, and reflux streams to be substantiallycontinuously introduced into reaction zone 28, while the gas and slurryeffluent streams are substantially continuously withdrawn from reactionzone 28. As used herein, the term “substantially continuously” shallmean for a period of at least 10 hours interrupted by less than 10minutes. During oxidation, it is preferred for the oxidizable compound(e.g., para-xylene) to be substantially continuously introduced intoreaction zone 28 at a rate of at least about 8,000 kilograms per hour,more preferably at a rate in the range of from about 15,000 to about200,000 kilograms per hour, still more preferably in the range of fromabout 22,000 to about 150,000 kilograms per hour, and most preferably inthe range of from 30,000 to 100,000 kilograms per hour. Although it isgenerally preferred for the flow rates of the incoming feed, oxidant,and reflux streams to be substantially steady, it is now noted that oneembodiment of the presenting invention contemplates pulsing the incomingfeed, oxidant, and/or reflux stream in order to improve mixing and masstransfer. When the incoming feed, oxidant, and/or reflux stream areintroduced in a pulsed fashion, it is preferred for their flow rates tovary within about 0 to about 500 percent of the steady-state flow ratesrecited herein, more preferably within about 30 to about 200 percent ofthe steady-state flow rates recited herein, and most preferably within80 to 120 percent of the steady-state flow rates recited herein.

The average space-time rate of reaction (STR) in bubble column oxidationreactor 20 is defined as the mass of the oxidizable compound fed perunit volume of reaction medium 36 per unit time (e.g., kilograms ofpara-xylene fed per cubic meter per hour). In conventional usage, theamount of oxidizable compound not converted to product would typicallybe subtracted from the amount of oxidizable compound in the feed streambefore calculating the STR. However, conversions and yields aretypically high for many of the oxidizable compounds preferred herein(e.g., para-xylene), and it is convenient to define the term herein asstated above. For reasons of capital cost and operating inventory, amongothers, it is generally preferred that the reaction be conducted with ahigh STR. However, conducting the reaction at increasingly higher STRmay affect the quality or yield of the partial oxidation. Bubble columnreactor 20 is particularly useful when the STR of the oxidizablecompound (e.g., para-xylene) is in the range of from about 25 kilogramsper cubic meter per hour to about 400 kilograms per cubic meter perhour, more preferably in the range of from about 30 kilograms per cubicmeter per hour to about 250 kilograms per cubic meter per hour, stillmore preferably from about 35 kilograms per cubic meter per hour toabout 150 kilograms per cubic meter per hour, and most preferably in therange of from 40 kilograms per cubic meter per hour to 100 kilograms percubic meter per hour.

The oxygen-STR in bubble column oxidation reactor 20 is defined as theweight of molecular oxygen consumed per unit volume of reaction medium36 per unit time (e.g., kilograms of molecular oxygen consumed per cubicmeter per hour). For reasons of capital cost and oxidative consumptionof solvent, among others, it is generally preferred that the reaction beconducted with a high oxygen-STR. However, conducting the reaction atincreasingly higher oxygen-STR eventually reduces the quality or yieldof the partial oxidation. Without being bound by theory, it appears thatthis possibly relates to the transfer rate of molecular oxygen from thegas phase into the liquid at the interfacial surface area and thenceinto the bulk liquid. Too high an oxygen-STR possibly leads to too low adissolved oxygen content in the bulk liquid phase of the reactionmedium.

The global-average-oxygen-STR is defined herein as the weight of alloxygen consumed in the entire volume of reaction medium 36 per unit time(e.g., kilograms of molecular oxygen consumed per cubic meter per hour).Bubble column reactor 20 is particularly useful when theglobal-average-oxygen-STR is in the range of from about 25 kilograms percubic meter per hour to about 400 kilograms per cubic meter per hour,more preferably in the range of from about 30 kilograms per cubic meterper hour to about 250 kilograms per cubic meter per hour, still morepreferably from about 35 kilograms per cubic meter per hour to about 150kilograms per cubic meter per hour, and most preferably in the range offrom 40 kilograms per cubic meter per hour to 100 kilograms per cubicmeter per hour.

During oxidation in bubble column reactor 20, it is preferred for theratio of the mass flow rate of the total solvent (from both the feed andreflux streams) to the mass flow rate of the oxidizable compoundentering reaction zone 28 to be maintained in the range of from about2:1 to about 50:1, more preferably in the range of from about 5:1 toabout 40:1, and most preferably in the range of from 7.5:1 to 25:1.Preferably, the ratio of the mass flow rate of solvent introduced aspart of the feed stream to the mass flow rate of solvent introduced aspart of the reflux stream is maintained in the range of from about 0.5:1to no reflux stream flow whatsoever, more preferably in the range offrom about 0.5:1 to about 4:1, still more preferably in the range offrom about 1:1 to about 2:1, and most preferably in the range of from1.25:1 to 1.5:1.

During liquid-phase oxidation in bubble column reactor 20, it ispreferred for the oxidant stream to be introduced into bubble columnreactor 20 in an amount that provides molecular oxygen somewhatexceeding the stoichiometric oxygen demand. The amount of excessmolecular oxygen required for best results with a particular oxidizablecompound affects the overall economics of the liquid-phase oxidation.During liquid-phase oxidation in bubble column reactor 20, it ispreferred that the ratio of the mass flow rate of the oxidant stream tothe mass flow rate of the oxidizable organic compound (e.g.,para-xylene) entering reactor 20 is maintained in the range of fromabout 0.5:1 to about 20:1, more preferably in the range of from about1:1 to about 10:1, and most preferably in the range of from 2:1 to 6:1.

Referring again to FIG. 1, the feed, oxidant, and reflux streamsintroduced into bubble column reactor 20 cooperatively form at least aportion of multi-phase reaction medium 36. Reaction medium 36 ispreferably a three-phase medium comprising a solid phase, a liquidphase, and a gas phase. As mentioned above, oxidation of the oxidizablecompound (e.g., para-xylene) takes place predominately in the liquidphase of reaction medium 36. Thus, the liquid phase of reaction medium36 comprises dissolved oxygen and the oxidizable compound. Theexothermic nature of the oxidation reaction that takes place in bubblecolumn reactor 20 causes a portion of the solvent (e.g., acetic acid andwater) introduced via feed inlets 32 a,b,c,d to boil/vaporize. Thus, thegas phase of reaction medium 36 in reactor 20 is formed primarily ofvaporized solvent and an undissolved, unreacted portion of the oxidantstream.

Certain prior art oxidation reactors employ heat exchange tubes/fins toheat or cool the reaction medium. However, such heat exchange structuresmay be undesirable in the inventive reactor and process describedherein. Thus, it is preferred for bubble column reactor 20 to includesubstantially no surfaces that contact reaction medium 36 and exhibit atime-averaged heat flux greater than 30,000 watts per meter squared. Inaddition, it is preferred for less than about 50 percent of thetime-averaged heat of reaction of reaction medium 36 to be removed byheat exchange surfaces, more preferably less than about 30 percent ofthe heat of reaction is removed by heat exchange surfaces, and mostpreferably less than 10 percent of the heat or reaction is removed byheat exchange surfaces.

The concentration of dissolved oxygen in the liquid phase of reactionmedium 36 is a dynamic balance between the rate of mass transfer fromthe gas phase and the rate of reactive consumption within the liquidphase (i.e. it is not set simply by the partial pressure of molecularoxygen in the supplying gas phase, though this is one factor in thesupply rate of dissolved oxygen and it does affect the limiting upperconcentration of dissolved oxygen). The amount of dissolved oxygenvaries locally, being higher near bubble interfaces. Globally, theamount of dissolved oxygen depends on the balance of supply and demandfactors in different regions of reaction medium 36. Temporally, theamount of dissolved oxygen depends on the uniformity of gas and liquidmixing relative to chemical consumption rates. In designing to matchappropriately the supply of and demand for dissolved oxygen in theliquid phase of reaction medium 36, it is preferred for thetime-averaged and volume-averaged oxygen concentration in the liquidphase of reaction medium 36 to be maintained above about 1 ppm molar,more preferably in the range from about 4 to about 1,000 ppm molar,still more preferably in the range from about 8 to about 500 ppm molar,and most preferably in the range from 12 to 120 ppm molar.

The liquid-phase oxidation reaction carried out in bubble column reactor20 is preferably a precipitating reaction that generates solids. Morepreferably, the liquid-phase oxidation carried out in bubble columnreactor 20 causes at least about 10 weight percent of the oxidizablecompound (e.g., para-xylene) introduced into reaction zone 28 to form asolid compound (e.g., crude terephthalic acid particles) in reactionmedium 36. Still more preferably, the liquid-phase oxidation causes atleast about 50 weight percent of the oxidizable compound to form a solidcompound in reaction medium 36. Most preferably, the liquid-phaseoxidation causes at least 90 weight percent of the oxidizable compoundto form a solid compound in reaction medium 36. It is preferred for thetotal amount of solids in reaction medium 36 to be greater than about 3percent by weight on a time-averaged and volume-averaged basis. Morepreferably, the total amount of solids in reaction medium 36 ismaintained in the range of from about 5 to about 40 weight percent,still more preferably in the range of from about 10 to about 35 weightpercent, and most preferably in the range of from 15 to 30 weightpercent. It is preferred for a substantial portion of the oxidationproduct (e.g., terephthalic acid) produced in bubble column reactor 20to be present in reaction medium 36 as solids, as opposed to remainingdissolved in the liquid phase of reaction medium 36. The amount of thesolid phase oxidation product present in reaction medium 36 ispreferably at least about 25 percent by weight of the total oxidationproduct (solid and liquid phase) in reaction medium 36, more preferablyat least about 75 percent by weight of the total oxidation product inreaction medium 36, and most preferably at least 95 percent by weight ofthe total oxidation product in reaction medium 36. The numerical rangesprovided above for the amount of solids in reaction medium 36 apply tosubstantially steady-state operation of bubble column 20 over asubstantially continuous period of time, not to start-up, shut-down, orsub-optimal operation of bubble column reactor 20. The amount of solidsin reaction medium 36 is determined by a gravimetric method. In thisgravimetric method, a representative portion of slurry is withdrawn fromthe reaction medium and weighed. At conditions that effectively maintainthe overall solid-liquid partitioning present within the reactionmedium, free liquid is removed from the solids portion by sedimentationor filtration, effectively without loss of precipitated solids and withless than about 10 percent of the initial liquid mass remaining with theportion of solids. The remaining liquid on the solids is evaporated todryness, effectively without sublimation of solids. The remainingportion of solids is weighed. The ratio of the weight of the portion ofsolids to the weight of the original portion of slurry is the fractionof solids, typically expressed as a percentage.

The precipitating reaction carried out in bubble column reactor 20 cancause fouling (i.e., solids build-up) on the surface of certain rigidstructures that contact reaction medium 36. Thus, in one embodiment ofthe present invention, it is preferred for bubble column reactor 20 toinclude substantially no internal heat exchange, stirring, or bafflingstructures in reaction zone 28 because such structures would be prone tofouling. If internal structures are present in reaction zone 28, it isdesirable to avoid internal structures having outer surfaces thatinclude a significant amount of upwardly facing planar surface areabecause such upwardly facing planar surfaces would be highly prone tofouling. Thus, if any internal structures are present in reaction zone28, it is preferred for less than about 20 percent of the total upwardlyfacing exposed outer surface area of such internal structures to beformed by substantially planar surfaces inclined less than about 15degrees from horizontal. Internal structures with this type ofconfiguration are referred to herein as having a “non-fouling”configuration.

Referring again to FIG. 1, the physical configuration of bubble columnreactor 20 helps provide for optimized oxidation of the oxidizablecompound (e.g., para-xylene) with minimal impurity generation. It ispreferred for elongated reaction section 24 of vessel shell 22 toinclude a substantially cylindrical main body 46 and a lower head 48.The upper end of reaction zone 28 is defined by a horizontal plane 50extending across the top of cylindrical main body 46. A lower end 52 ofreaction zone 28 is defined by the lowest internal surface of lower head48. Typically, lower end 52 of reaction zone 28 is located proximate theopening for slurry outlet 38. Thus, elongated reaction zone 28 definedwithin bubble column reactor 20 has a maximum length “L” measured fromthe top end 50 to the bottom end 52 of reaction zone 28 along the axisof elongation of cylindrical main body 46. The length “L” of reactionzone 28 is preferably in the range of from about 10 to about 100 meters,more preferably in the range of from about 20 to about 75 meters, andmost preferably in the range of from 25 to 50 meters. Reaction zone 28has a maximum diameter (width) “D” that is typically equal to themaximum internal diameter of cylindrical main body 46. The maximumdiameter “D” of reaction zone 28 is preferably in the range of fromabout 1 to about 12 meters, more preferably in the range of from about 2to about 10 meters, still more preferably in the range of from about 3.1to about 9 meters, and most preferably in the range of from 4 to 8meters. In a preferred embodiment of the present invention, reactionzone 28 has a length-to-diameter “L:D” ratio in the range of from about6:1 to about 30:1. Still more preferably, reaction zone 28 has an L:Dratio in the range of from about 8:1 to about 20:1. Most preferably,reaction zone 28 has an L:D ratio in the range of from 9:1 to 15:1.

As discussed above, reaction zone 28 of bubble column reactor 20receives multi-phase reaction medium 36. Reaction medium 36 has a bottomend coincident with lower end 52 of reaction zone 28 and a top endlocated at upper surface 44. Upper surface 44 of reaction medium 36 isdefined along a horizontal plane that cuts through reaction zone 28 at avertical location where the contents of reaction zone 28 transitionsfrom a gas-phase-continuous state to a liquid-phase-continuous state.Upper surface 44 is preferably positioned at the vertical location wherethe local time-averaged gas hold-up of a thin horizontal slice of thecontents of reaction zone 28 is 0.9.

Reaction medium 36 has a maximum height “H” measured between its upperand lower ends. The maximum width “W” of reaction medium 36 is typicallyequal to the maximum diameter “D” of cylindrical main body 46. Duringliquid-phase oxidation in bubble column reactor 20, it is preferred thatH is maintained at about 60 to about 120 percent of L, more preferablyabout 80 to about 110 percent of L, and most preferably 85 to 100percent of L. In a preferred embodiment of the present invention,reaction medium 36 has a height-to-width “H:W” ratio greater than about3:1. More preferably, reaction medium 36 has an H:W ratio in the rangeof from about 7:1 to about 25:1. Still more preferably, reaction medium36 has an H:W ratio in the range of from about 8:1 to about 20:1. Mostpreferably, reaction medium 36 has an H:W ratio in the range of from 9:1to 15:1. In one embodiment of the invention, L=H and D=W so that variousdimensions or ratios provide herein for L and D also apply to H and W,and vice-versa.

The relatively high L:D and H:W ratios provided in accordance with anembodiment of the invention can contribute to several importantadvantages of the inventive system. As discussed in further detailbelow, it has been discovered that higher L:D and H:W ratios, as well ascertain other features discussed below, can promote beneficial verticalgradients in the concentrations of molecular oxygen and/or theoxidizable compound (e.g., para-xylene) in reaction medium 36. Contraryto conventional wisdom, which would favor a well-mixed reaction mediumwith relatively uniform concentrations throughout, it has beendiscovered that the vertical staging of the oxygen and/or the oxidizablecompound concentrations facilitates a more effective and economicaloxidation reaction. Minimizing the oxygen and oxidizable compoundconcentrations near the top of reaction medium 36 can help avoid loss ofunreacted oxygen and unreacted oxidizable compound through upper gasoutlet 40. However, if the concentrations of oxidizable compound andunreacted oxygen are low throughout reaction medium 36, then the rateand/or selectivity of oxidation are reduced. Thus, it is preferred forthe concentrations of molecular oxygen and/or the oxidizable compound tobe significantly higher near the bottom of reaction medium 36 than nearthe top of reaction medium 36.

In addition, high L:D and H:W ratios cause the pressure at the bottom ofreaction medium 36 to be substantially greater than the pressure at thetop of reaction medium 36. This vertical pressure gradient is a resultof the height and density of reaction medium 36. One advantage of thisvertical pressure gradient is that the elevated pressure at the bottomof the vessel drives more oxygen solubility and mass transfer than wouldotherwise be achievable at comparable temperatures and overheadpressures in shallow reactors. Thus, the oxidation reaction can becarried out at lower temperatures than would be required in a shallowervessel. When bubble column reactor 20 is used for the partial oxidationof para-xylene to crude terephthalic acid (CTA), the ability to operateat lower reaction temperatures with the same or better oxygen masstransfer rates has a number of advantages. For example, low temperatureoxidation of para-xylene reduces the amount of solvent burned during thereaction. As discussed in further detail below, low temperatureoxidation also favors the formation of small, high surface area, looselybound, easily dissolved CTA particles, which can be subjected to moreeconomical purification techniques than the large, low surface area,dense CTA particles produced by conventional high temperature oxidationprocesses.

During oxidation in reactor 20, it is preferred for the time-averagedand volume-averaged temperature of reaction medium 36 to be maintainedin the range of from about 125 to about 200° C., more preferably in therange of from about 140 to about 180° C., and most preferably in therange of from 150 to 170° C. The overhead pressure above reaction medium36 is preferably maintained in the range of from about 1 to about 20 bargauge (barg), more preferably in the range of from about 2 to about 12barg, and most preferably in the range of from 4 to 8 barg. Preferably,the pressure difference between the top of reaction medium 36 and thebottom of reaction medium 36 is in the range of from about 0.4 to about5 bar, more preferably the pressure difference is in the range of fromabout 0.7 to about 3 bars, and most preferably the pressure differenceis 1 to 2 bar. Although it is generally preferred for the overheadpressure above reaction medium 36 to be maintained at a relativelyconstant value, one embodiment of the present invention contemplatespulsing the overhead pressure to facilitate improved mixing and/or masstransfer in reaction medium 36. When the overhead pressure is pulsed, itis preferred for the pulsed pressures to range between about 60 to about140 percent of the steady-state overhead pressure recited herein, morepreferably between about 85 and about 115 percent of the steady-stateoverhead pressure recited herein, and most preferably between 95 and 105percent of the steady-state overhead pressure recited herein.

A further advantage of the high L:D ratio of reaction zone 28 is that itcan contribute to an increase in the average superficial velocity ofreaction medium 36. The term “superficial velocity” and “superficial gasvelocity,” as used herein with reference to reaction medium 36, shalldenote the volumetric flow rate of the gas phase of reaction medium 36at an elevation in the reactor divided by the horizontal cross-sectionalarea of the reactor at that elevation. The increased superficialvelocity provided by the high L:D ratio of reaction zone 28 can promotelocal mixing and increase the gas hold-up of reaction medium 36. Thetime-averaged superficial velocities of reaction medium 36 atone-quarter height, half height, and/or three-quarter height of reactionmedium 36 are preferably greater than about 0.3 meters per second, morepreferably in the range of from about 0.8 to about 5 meters per second,still more preferably in the range of from about 0.9 to about 4 metersper second, and most preferably in the range of from 1 to 3 meters persecond.

Referring again to FIG. 1, disengagement section 26 of bubble columnreactor 20 is simply a widened portion of vessel shell 22 locatedimmediately above reaction section 24. Disengagement section 26 reducesthe velocity of the upwardly-flowing gas phase in bubble column reactor20 as the gas phase rises above the upper surface 44 of reaction medium36 and approaches gas outlet 40. This reduction in the upward velocityof the gas phase helps facilitate removal of entrained liquids and/orsolids in the upwardly flowing gas phase and thereby reduces undesirableloss of certain components present in the liquid phase of reactionmedium 36.

Disengagement section 26 preferably includes a generally frustoconicaltransition wall 54, a generally cylindrical broad sidewall 56, and anupper head 58. The narrow lower end of transition wall 54 is coupled tothe top of cylindrical main body 46 of reaction section 24. The wideupper end of transition wall 54 is coupled to the bottom of broadsidewall 56. It is preferred for transition wall 54 to extend upwardlyand outwardly from its narrow lower end at an angle in the range of fromabout 10 to about 70 degrees from vertical, more preferably in the rangeof about 15 to about 50 degrees from vertical, and most preferably inthe range of from 15 to 45 degrees from vertical. Broad sidewall 56 hasa maximum diameter “X” that is generally greater than the maximumdiameter “D” of reaction section 24, though when the upper portion ofreaction section 24 has a smaller diameter than the overall maximumdiameter of reaction section 24, then X may actually be smaller than D.In a preferred embodiment of the present invention, the ratio of thediameter of broad sidewall 56 to the maximum diameter of reactionsection 24 “X:D” is in the range of from about 0.8:1 to about 4:1, mostpreferably in the range of from 1.1:1 to 2:1. Upper head 58 is coupledto the top of broad sidewall 56. Upper head 58 is preferably a generallyelliptical head member defining a central opening that permits gas toescape disengagement zone 30 via gas outlet 40. Alternatively, upperhead 58 may be of any shape, including conical. Disengagement zone 30has a maximum height “Y” measured from the top 50 of reaction zone 28 tothe upper most portion of disengagement zone 30. The ratio of the lengthof reaction zone 28 to the height of disengagement zone 30 “L:Y” ispreferably in the range of from about 2:1 to about 24:1, more preferablyin the range of from about 3:1 to about 20:1, and most preferably in therange of from 4:1 to 16:1.

Referring now to FIGS. 1-5, the location and configuration of oxidantsparger 34 will now be discussed in greater detail. FIGS. 2 and 3 showthat oxidant sparger 34 can include a ring member 60 and a pair ofoxidant entry conduits 64 a,b. Conveniently, these oxidant entryconduits 64 a,b can enter the vessel at an elevation above the ringmember 60 and then turn downwards as shown in FIG. 2. Alternatively, anoxidant entry conduit may enter the vessel below the ring member 60 oron about the same horizontal plane as ring member 60. Each oxidant entryconduit 64 a,b includes a first end coupled to a respective oxidantinlet 66 a,b formed in the vessel shell 22 and a second end fluidlycoupled to ring member 60. Ring member 60 is preferably formed ofconduits, more preferably of a plurality of straight conduit sections,and most preferably a plurality of straight pipe sections, rigidlycoupled to one another to form a tubular polygonal ring. Preferably,ring member 60 is formed of at least 3 straight pipe sections, morepreferably 6 to 10 pipe sections, and most preferably 8 pipe sections.Accordingly, when ring member 60 is formed of 8 pipe sections, it has agenerally octagonal configuration. It is preferred for the pipe sectionsthat make up oxidant entry conduits 64a,b and ring member 60 to have anominal diameter greater than about 0.1 meter, more preferable in therange of from about 0.2 to about 2 meters, and most preferably in therange of from 0.25 to 1 meters. As perhaps best illustrated in FIG. 3,it is preferred that substantially no openings are formed in the upperportion of sparger ring 60.

As perhaps best illustrated in FIGS. 4 and 5, the bottom portion ofoxidant sparger ring 60 presents a plurality of oxidant openings 68.Oxidant openings 68 are preferably configured such that at least about 1percent of the total open area defined by oxidant openings 68 is locatedbelow the centerline 64 (FIG. 5) of ring member 60, where centerline 64is located at the elevation of the volumetric centroid of ring member60. More preferably, at least about 5 percent of the total open areadefined by all oxidant openings 68 is located below centerline 64, withat least about 2 percent of the total open area being defined byopenings 68 that discharge the oxidant stream in a generally downwarddirection within about 30 degrees of vertical. Still more preferably, atleast about 20 percent of the total open area defined by all oxidantopenings 68 is located below centerline 64, with at least about 10percent of the total open area being defined by openings 68 thatdischarge the oxidant stream in a generally downward direction within 30degrees of vertical. Most preferably, at least about 75 percent of thetotal open area defined by all oxidant openings 68 is located belowcenterline 64, with at least about 40 percent of the total open areabeing defined by openings 68 that discharge the oxidant stream in agenerally downward direction within 30 degrees of vertical. The fractionof the total open area defined by all oxidant openings 68 that arelocated above centerline 64 is preferably less than about 75 percent,more preferably less than about 50 percent, still more preferably lessthan about 25 percent, and most preferably less than 5 percent.

As illustrated in FIGS. 4 and 5, oxidant openings 68 include downwardopenings 68 a and skewed openings 68 b. Downward openings 68 a areconfigured to discharge the oxidant stream generally downwardly at anangle within about 30 degrees of vertical, more preferably within about15 degrees of vertical, and most preferably within 5 degrees ofvertical. Referring now to FIG. 5, skewed openings 68 b are configuredto discharge the oxidant stream generally outwardly and downwardly at anangle “A” that is in the range of from about 15 to about 75 degrees fromvertical, more preferably angle A is in the range of from about 30 toabout 60 degrees from vertical, and most preferably angle A is in therange of from 40 to 50 degrees from vertical.

It is preferred for substantially all oxidant openings 68 to haveapproximately the same diameter. The diameter of oxidant openings 68 ispreferably in the range of from about 2 to about 300 millimeters, morepreferably in the range of from about 4 to about 120 millimeters, andmost preferably in the range of from 8 to 60 millimeters. The totalnumber of oxidant openings 68 in ring member 60 is selected to meet thelow pressure drop criteria detailed below. Preferably, the total numberof oxidant openings 68 formed in ring member 60 is at least about 10,more preferably the total number of oxidant openings 68 is in the rangeof from about 20 to about 200, and most preferably the total number ofoxidant openings 68 is in the range of from 40 to 100.

Although FIGS. 1-5 illustrate a very specific configuration for oxidantsparger 34, it is now noted that a variety of oxidant spargerconfigurations can be employed to achieve the advantages describedherein. For example, an oxidant sparger does not necessarily need tohave the octagonal ring member configuration illustrated in FIGS. 1-5.Rather, it is possible for an oxidant sparger to be formed of anyconfiguration of flow conduit(s) that employs a plurality ofspaced-apart openings for discharging the oxidant stream. The size,number, and discharge direction of the oxidant openings in the flowconduit are preferably within the ranges stated above. Further, theoxidant sparger is preferably configured to provide the azimuthal andradial distribution of molecular oxygen described above.

Regardless of the specific configuration of oxidant sparger 34, it ispreferred for the oxidant sparger to be physically configured andoperated in a manner that minimizes the pressure drop associated withdischarging the oxidant stream out of the flow conduit(s), through theoxidant openings, and into the reaction zone. Such pressure drop iscalculated as the time-averaged static pressure of the oxidant streaminside the flow conduit at oxidant inlets 66 a,b of the oxidant spargerminus the time-averaged static pressure in the reaction zone at theelevation where one-half of the oxidant stream is introduced above thatvertical location and one-half of the oxidant stream is introduced belowthat vertical location. In a preferred embodiment of the presentinvention, the time-averaged pressure drop associated with dischargingthe oxidant stream from the oxidant sparger is less than about 0.3megapascal (MPa), more preferably less than about 0.2 MPa, still morepreferably less than about 0.1 MPa, and most preferably less than 0.05MPa.

Optionally, a continuous or intermittent flush can be provided tooxidant sparger 34 with a liquid (e.g., acetic acid, water, and/orpara-xylene) to prevent fouling of the oxidant sparger with solids. Whensuch a liquid flush is employed, it is preferred for an effective amountof the liquid (i.e., not just the minor amount of liquid droplets thatmight naturally be present in the oxidant stream) to be passed throughthe oxidant sparger and out of the oxidant openings for at least oneperiod of more than one minute each day. When a liquid is continuouslyor periodically discharged from oxidant sparger 34, it is preferred forthe time-averaged ratio of the mass flow rate of the liquid through theoxidant sparger to the mass flow rate of the molecular oxygen throughthe oxidant sparger to be in the range of from about 0.05:1 to about30:1, or in the range of from about 0.1:1 to about 2:1, or even in therange of from 0.2:1 to 1:1.

In many conventional bubble column reactors containing a multi-phasereaction medium, substantially all of the reaction medium located belowthe oxidant sparger (or other mechanism for introducing the oxidantstream into the reaction zone) has a very low gas hold-up value. Asknown in the art, “gas hold-up” is simply the volume fraction of amulti-phase medium that is in the gaseous state. Zones of low gashold-up in a medium can also be referred to as “unaerated” zones. Inmany conventional slurry bubble column reactors, a significant portionof the total volume of the reaction medium is located below the oxidantsparger (or other mechanism for introducing the oxidant stream into thereaction zone). Thus, a significant portion of the reaction mediumpresent at the bottom of conventional bubble column reactors isunaerated.

It has been discovered that minimizing the amount of unaerated zones ina reaction medium subjected to oxidization in a bubble column reactorcan minimize the generation of certain types of undesirable impurities.Unaerated zones of a reaction medium contain relatively few oxidantbubbles. This low volume of oxidant bubbles reduces the amount ofmolecular oxygen available for dissolution into the liquid phase of thereaction medium. Thus, the liquid phase in an unaerated zone of thereaction medium has a relatively low concentration of molecular oxygen.These oxygen-starved, unaerated zones of the reaction medium have atendency to promote undesirable side reactions, rather than the desiredoxidation reaction. For example, when para-xylene is partially oxidizedto form terephthalic acid, insufficient oxygen availability in theliquid phase of the reaction medium can cause the formation ofundesirably high quantities of benzoic acid and coupled aromatic rings,notably including highly undesirable colored molecules known asfluorenones and anthraquinones.

In accordance with one embodiment of the present invention, liquid-phaseoxidation is carried out in a bubble column reactor configured andoperated in a manner such that the volume fraction of the reactionmedium with low gas hold-up values is minimized This minimization ofunaerated zones can be quantified by theoretically partitioning theentire volume of the reaction medium into 2,000 discrete horizontalslices of uniform volume. With the exception of the highest and lowesthorizontal slices, each horizontal slice is a discrete volume bounded onits sides by the sidewall of the reactor and bounded on its top andbottom by imaginary horizontal planes. The highest horizontal slice isbounded on its bottom by an imaginary horizontal plane and on its top bythe upper surface of the reaction medium. The lowest horizontal slice isbounded on its top by an imaginary horizontal plane and on its bottom bythe lower end of the vessel. Once the reaction medium has beentheoretically partitioned into 2,000 discrete horizontal slices of equalvolume, the time-averaged and volume-averaged gas hold-up of eachhorizontal slice can be determined When this method of quantifying theamount of unaerated zones is employed, it is preferred for the number ofhorizontal slices having a time-averaged and volume-averaged gas hold-upless than 0.1 to be less than 30, more preferably less than 15, stillmore preferably less than 6, even more preferably less than 4, and mostpreferably less than 2. It is preferred for the number of horizontalslices having a gas hold-up less than 0.2 to be less than 80, morepreferably less than 40, still more preferably less than 20, even morepreferably less than 12, and most preferably less than 5. It ispreferred for the number of horizontal slices having a gas hold-up lessthan 0.3 to be less than 120, more preferably less than 80, still morepreferably less than 40, even more preferably less than 20, and mostpreferably less than 15.

Referring again to FIGS. 1 and 2, it has been discovered thatpositioning oxidant sparger 34 lower in reaction zone 28 providesseveral advantages, including reduction of the amount of unaerated zonesin reaction medium 36. Given a height “H” of reaction medium 36, alength “L” of reaction zone 28, and a maximum diameter “D” of reactionzone 28, it is preferred for a majority (i.e., >50 percent by weight) ofthe oxidant stream to be introduced into reaction zone 28 within about0.025 H, 0.022 L, and/or 0.25 D of lower end 52 of reaction zone 28.More preferably, a majority of the oxidant stream is introduced intoreaction zone 28 within about 0.02 H, 0.018 L, and/or 0.2 D of lower end52 of reaction zone 28. Most preferably, a majority of the oxidantstream is introduced into reaction zone 28 within 0.015 H, 0.013 L,and/or 0.15 D of lower end 52 of reaction zone 28.

In the embodiment illustrated in FIG. 2, the vertical distance “Y₁”between lower end 52 of reaction zone 28 and the outlet of upper oxidantopenings 68 of oxidant sparger 34 is less than about 0.25 H, 0.022 L,and/or 0.25 D, so that substantially all of the oxidant stream entersreaction zone 28 within about 0.25 H, 0.022 L, and/or 0.25 D of lowerend 52 of reaction zone 28. More preferably, Y₁ is less than about 0.02H, 0.018 L, and/or 0.2 D. Most preferably, Y₁ is less than 0.015 H,0.013 L, and/or 0.15 D, but more than 0.005 H, 0.004 L, and/or 0.06 D.FIG. 2 illustrates a tangent line 72 at the location where the bottomedge of cylindrical main body 46 of vessel shell 22 joins with the topedge of elliptical lower head 48 of vessel shell 22. Alternatively,lower head 48 can be of any shape, including conical, and the tangentline is still defined as the bottom edge of cylindrical main body 46.The vertical distance “Y₂” between tangent line 72 and the top ofoxidant sparger 34 is preferably at least about 0.0012 H, 0.001 L,and/or 0.01 D; more preferably at least about 0.005 H, 0.004 L, and/or0.05 D; and most preferably at least 0.O1 H, 0.008 L, and/or 0.1 D. Thevertical distance “Y₃” between lower end 52 of reaction zone 28 and theoutlet of lower oxidant openings 70 of oxidant sparger 34 is preferablyless than about 0.015 H, 0.013 L, and/or 0.15 D; more preferably lessthan about 0.012 H, 0.01 L, and/or 0.1 D; and most preferably less than0.O1 H, 0.008 L, and/or 0.075 D, but more than 0.003 H, 0.002 L, and/or0.025 D.

In addition to the advantages provided by minimizing unaerated zones(i.e., zones with low gas hold-up) in reaction medium 36, it has beendiscovered that oxidation can be enhanced by maximizing the gas hold-upof the entire reaction medium 36. Reaction medium 36 preferably hastime-averaged and volume-averaged gas hold-up of at least about 0.4,more preferably in the range of from about 0.6 to about 0.9, and mostpreferably in the range of from 0.65 to 0.85. Several physical andoperational attributes of bubble column reactor 20 contribute to thehigh gas hold-up discussed above. For example, for a given reactor sizeand flow of oxidant stream, the high L:D ratio of reaction zone 28yields a lower diameter which increases the superficial velocity inreaction medium 36 which in turn increases gas hold-up. Additionally,the actual diameter of a bubble column and the L:D ratio are known toinfluence the average gas hold-up even for a given constant superficialvelocity. In addition, the minimization of unaerated zones, particularlyin the bottom of reaction zone 28, contributes to an increased gashold-up value. Further, the overhead pressure and mechanicalconfiguration of the bubble column reactor can affect operatingstability at the high superficial velocities and gas hold-up valuesdisclosed herein.

Referring again to FIG. 1, it has been discovered that improveddistribution of the oxidizable compound (e.g., para-xylene) in reactionmedium 36 can be provided by introducing the liquid-phase feed streaminto reaction zone 28 at multiple vertically-spaced locations.Preferably, the liquid-phase feed stream is introduced into reactionzone 28 via at least 3 feed openings, more preferably at least 4 feedopenings. As used herein, the term “feed openings” shall denote openingswhere the liquid-phase feed stream is discharged into reaction zone 28for mixing with reaction medium 36. It is preferred for at least 2 ofthe feed openings to be vertically-spaced from one another by at leastabout 0.5 D, more preferably at least about 1.5 D, and most preferablyat least 3 D. However, it is preferred for the highest feed opening tobe vertically-spaced from the lowest oxidant opening by not more thanabout 0.75 H, 0.65 L, and/or 8 D; more preferably not more than about0.5 H, 0.4 L, and/or 5 D; and most preferably not more than 0.4 H, 0.35L, and/or 4 D.

Although it is desirable to introduce the liquid-phase feed stream atmultiple vertical locations, it has also been discovered that improveddistribution of the oxidizable compound in reaction medium 36 isprovided if the majority of the liquid-phase feed stream is introducedinto the bottom half of reaction medium 36 and/or reaction zone 28.Preferably, at least about 75 weight percent of the liquid-phase feedstream is introduced into the bottom half of reaction medium 36 and/orreaction zone 28. Most preferably, at least 90 weight percent of theliquid-phase feed stream is introduced into the bottom half of reactionmedium 36 and/or reaction zone 28. In addition, it is preferred for atleast about 30 weight percent of the liquid-phase feed stream to beintroduced into reaction zone 28 within about 1.5 D of the lowestvertical location where the oxidant stream is introduced into reactionzone 28. This lowest vertical location where the oxidant stream isintroduced into reaction zone 28 is typically at the bottom of oxidantsparger; however, a variety of alternative configurations forintroducing the oxidant stream into reaction zone 28 are contemplated bya preferred embodiment of the present invention. Preferably, at leastabout 50 weight percent of the liquid-phase feed is introduced withinabout 2.5 D of the lowest vertical location where the oxidant stream isintroduced into reaction zone 28. Preferably, at least about 75 weightpercent of the liquid-phase feed stream is introduced within about 5 Dof the lowest vertical location where the oxidant stream is introducedinto reaction zone 28.

Each feed opening defines an open area through which the feed isdischarged. It is preferred that at least about 30 percent of thecumulative open area of all the feed inlets is located within about 1.5D of the lowest vertical location where the oxidant stream is introducedinto reaction zone 28. Preferably, at least about 50 percent of thecumulative open area of all the feed inlets is located within about 2.5D of the lowest vertical location where the oxidant stream is introducedinto reaction zone 28. Preferably, at least about 75 percent of thecumulative open area of all the feed inlets is located within about 5 Dof the lowest vertical location where the oxidant stream is introducedinto reaction zone 28.

Referring again to FIG. 1, in one embodiment of the present invention,feed inlets 32 a,b,c,d are simply a series of vertically-alignedopenings along one side of vessel shell 22. These feed openingspreferably have substantially similar diameters of less than about 7centimeters, more preferably in the range of from about 0.25 to about 5centimeters, and most preferably in the range of from 0.4 to 2centimeters. Bubble column reactor 20 is preferably equipped with asystem for controlling the flow rate of the liquid-phase feed stream outof each feed opening. Such flow control system preferably includes anindividual flow control valve 74 a,b,c,d for each respective feed inlet32 a,b,c,d. In addition, it is preferred for bubble column reactor 20 tobe equipped with a flow control system that allows at least a portion ofthe liquid-phase feed stream to be introduced into reaction zone 28 atan elevated inlet superficial velocity of at least about 2 meters persecond, more preferably at least about 5 meters per second, still morepreferably at least about 6 meters per second, and most preferably inthe range of from 8 to 20 meters per second. As used herein, the term“inlet superficial velocity” denotes the time-averaged volumetric flowrate of the feed stream out of the feed opening divided by the area ofthe feed opening. Preferably, at least about 50 weight percent of thefeed stream is introduced into reaction zone 28 at an elevated inletsuperficial velocity. Most preferably, substantially all the feed streamis introduced into reaction zone 28 at an elevated inlet superficialvelocity.

Referring now to FIGS. 6 and 7, an alternative system for introducingthe liquid-phase feed stream into reaction zone 28 is illustrated. Inthis embodiment, the feed stream is introduced into reaction zone 28 atfour different elevations. Each elevation is equipped with a respectivefeed distribution system 76 a,b,c,d. Each feed distribution system 76includes a main feed conduit 78 and a manifold 80. Each manifold 80 isprovided with at least two outlets 82,84 coupled to respective insertconduits 86,88, which extend into reaction zone 28 of vessel shell 22.Each insert conduit 86,88 presents a respective feed opening 87,89 fordischarging the feed stream into reaction zone 28. Feed openings 87,89preferably have substantially similar diameters of less than about 7centimeters, more preferably in the range of from about 0.25 to about 5centimeters, and most preferably in the range of from 0.4 to 2centimeters. It is preferred for feed openings 87,89 of each feeddistribution system 76 a,b,c,d to be diametrically opposed so as tointroduce the feed stream into reaction zone 28 in opposite directions.Further, it is preferred for the diametrically opposed feed openings86,88 of adjacent feed distribution systems 76 to be oriented at 90degrees of rotation relative to one another. In operation, theliquid-phase feed stream is charged to main feed conduit 78 andsubsequently enters manifold 80. Manifold 80 distributes the feed streamevenly for simultaneous introduction on opposite sides of reactor 20 viafeed openings 87,89.

FIG. 8 illustrates an alternative configuration wherein each feeddistribution system 76 is equipped with bayonet tubes 90,92 rather thaninsert conduits 86,88 (shown in FIG. 7). Bayonet tubes 90,92 projectinto reaction zone 28 and include a plurality of small feed openings94,96 for discharging the liquid-phase feed into reaction zone 28. It ispreferred for the small feed openings 94,96 of bayonet tubes 90,92 tohave substantially the same diameters of less than about 50 millimeters,more preferably about 2 to about 25 millimeters, and most preferably 4to 15 millimeters.

FIGS. 9-11 illustrate an alternative feed distribution system 100. Feeddistribution system 100 introduces the liquid-phase feed stream at aplurality of vertically-spaced and laterally-spaced locations withoutrequiring multiple penetrations of the sidewall of bubble column reactor20. Feed introduction system 100 generally includes a single inletconduit 102, a header 104, a plurality of upright distribution tubes106, a lateral support mechanism 108, and a vertical support mechanism110. Inlet conduit 102 penetrates the sidewall of main body 46 of vesselshell 22. Inlet conduit 102 is fluidly coupled to header 104. Header 104distributes the feed stream received from inlet conduit 102 evenly amongupright distribution tubes 106. Each distribution tube 106 has aplurality of vertically-spaced feed openings 112 a,b,c,d for dischargingthe feed stream into reaction zone 28. Lateral support mechanism 108 iscoupled to each distribution tube 106 and inhibits relative lateralmovement of distribution tubes 106. Vertical support mechanism 110 ispreferably coupled to lateral support mechanism 108 and to the top ofoxidant sparger 34. Vertical support mechanism 110 substantiallyinhibits vertical movement of distribution tubes 106 in reaction zone28. It is preferred for feed openings 112 to have substantially the samediameters of less than about 50 millimeters, more preferably about 2 toabout 25 millimeters, and most preferably 4 to 15 millimeters. Thevertical spacing of feed openings 112 of feed distribution system 100illustrated in FIGS. 9-11 can be substantially the same as describedabove with reference to the feed distribution system of FIG. 1.Optionally, feed openings can be elongated nozzles rather than simpleholes. Optionally, one or more flow deflection apparatus can lie outsideof the flow conduit and in path of fluids exiting therefrom into thereaction medium. Optionally, an opening near the bottom of a flowconduit can be sized to purge solids from inside the liquid-phase feeddistribution system, either continuously or intermittently. Optionally,mechanical devices such as flapper assemblies, check valves, excess flowvalves, power operated valves and the like may be used either to preventingress of solids during operational upsets or to discharge accumulatedsolids from within the liquid-phase feed distribution system.

It has been discovered that the flow patterns of the reaction medium inmany bubble column reactors can permit uneven azimuthal distribution ofthe oxidizable compound in the reaction medium, especially when theoxidizable compound is primarily introduced along one side of thereaction medium. As used herein, the term “azimuthal” shall denote anangle or spacing around the upright axis of elongation of the reactionzone. As used herein, “upright” shall mean within 45° of vertical. Inone embodiment of the present invention, the feed stream containing theoxidizable compound (e.g., para-xylene) is introduced into the reactionzone via a plurality of azimuthally-spaced feed openings. Theseazimuthally-spaced feed openings can help prevent regions of excessivelyhigh and excessively low oxidizable compound concentrations in thereaction medium. The various feed introduction systems illustrated inFIGS. 6-11 are examples of systems that provide proper azimuthal spacingof feed openings.

Referring again to FIG. 7, in order to quantify the azimuthally-spacedintroduction of the liquid-phase feed stream into the reaction medium,the reaction medium can be theoretically partitioned into four uprightazimuthal quadrants “Q₁,Q₂,Q₃,Q₄” of approximately equal volume. Theseazimuthal quadrants “Q₁,Q₂,Q₃,Q₄” are defined by a pair of imaginaryintersecting perpendicular vertical planes “P₁,P₂” extending beyond themaximum vertical dimension and maximum radial dimension of the reactionmedium. When the reaction medium is contained in a cylindrical vessel,the line of intersection of the imaginary intersecting vertical planesP₁,P₂ will be approximately coincident with the vertical centerline ofthe cylinder, and each azimuthal quadrant Q₁,Q₂,Q₃,Q₄ will be agenerally wedge-shaped vertical volume having a height equal to theheight of the reaction medium. It is preferred for a substantial portionof the oxidizable compound to be discharged into the reaction medium viafeed openings located in at least two different azimuthal quadrants.

In a preferred embodiment of the present invention, not more than about80 weight percent of the oxidizable compound is discharged into thereaction medium through feed openings that can be located in a singleazimuthal quadrant. More preferably, not more than about 60 weightpercent of the oxidizable compound is discharged into the reactionmedium through feed openings that can be located in a single azimuthalquadrant. Most preferably, not more than 40 weight percent of theoxidizable compound is discharged into the reaction medium through feedopenings that can be located in a single azimuthal quadrant. Theseparameters for azimuthal distribution of the oxidizable compound aremeasured when the azimuthal quadrants are azimuthally oriented such thatthe maximum possible amount of oxidizable compound is being dischargedinto one of the azimuthal quadrants. For example, if the entire feedstream is discharged into the reaction medium via two feed openings thatare azimuthally spaced from one another by 89 degrees, for purposes ofdetermining azimuthal distribution in four azimuthal quadrants, 100weight percent of the feed stream is discharged into the reaction mediumin a single azimuthal quadrant because the azimuthal quadrants can beazimuthally oriented in such a manner that both of the feed openings arelocated in a single azimuthal quadrant.

In addition to the advantages associated with the properazimuthal-spacing of the feed openings, it has also been discovered thatproper radial spacing of the feed openings in a bubble column reactorcan also be important. It is preferred for a substantial portion of theoxidizable compound introduced into the reaction medium to be dischargedvia feed openings that are radially spaced inwardly from the sidewall ofthe vessel. Thus, in one embodiment of the present invention, asubstantial portion of the oxidizable compound enters the reaction zonevia feed openings located in a “preferred radial feed zone” that isspaced inwardly from the upright sidewalls defining the reaction zone.

Referring again to FIG. 7, the preferred radial feed zone “FZ” can takethe shape of a theoretical upright cylinder centered in reaction zone 28and having an outer diameter “D_(O)” of 0.9 D, where “D” is the diameterof reaction zone 28. Thus, an outer annulus “OA” having a thickness of0.05 D is defined between the preferred radial feed zone FZ and theinside of the sidewall defining reaction zone 28. It is preferred forlittle or none of the oxidizable compound to be introduced into reactionzone 28 via feed openings located in this outer annulus OA.

In another embodiment, it is preferred for little or none of theoxidizable compound to be introduced into the center of reaction zone28. Thus, as illustrated in FIG. 8, the preferred radial feed zone FZcan take the shape of a theoretical upright annulus centered in reactionzone 28, having an outer diameter D_(O) of 0.9 D, and having an innerdiameter D_(I) of 0.2 D. Thus, in this embodiment, an inner cylinder IChaving a diameter of 0.2 D is “cut out” of the center of the preferredradial feed zone FZ. It is preferred for little or none of theoxidizable compound to be introduced into reaction zone 28 via feedopenings located in this inner cylinder IC.

In a preferred embodiment of the present invention, a substantialportion of the oxidizable compound is introduced into reaction medium 36via feed openings located in the preferred radial feed zone, regardlessof whether the preferred radial feed zone has the cylindrical or annularshape described above. More preferably, at least about 25 weight percentof the oxidizable compound is discharged into reaction medium 36 viafeed openings located in the preferred radial feed zone. Still morepreferably, at least about 50 weight percent of the oxidizable compoundis discharged into reaction medium 36 via feed openings located in thepreferred radial feed zone. Most preferably, at least 75 weight percentof the oxidizable compound is discharged into reaction medium 36 viafeed openings located in the preferred radial feed zone.

Although the theoretical azimuthal quadrants and theoretical preferredradial feed zone illustrated in FIGS. 7 and 8 are described withreference to the distribution of the liquid-phase feed stream, it hasbeen discovered that proper azimuthal and radial distribution of thegas-phase oxidant stream can also provide certain advantages. Thus, inone embodiment of the present invention, the description of theazimuthal and radial distribution of the liquid-phase feed stream,provided above, also applies to the manner in which the gas-phaseoxidant stream is introduced into the reaction medium 36.

Referring now to FIGS. 12 and 13, there is illustrated an alternativebubble column reactor 200 having a reactor-in-reactor configuration.Bubble column reactor 200 includes an external reactor 202 and aninternal reactor 204, with internal reactor 204 being at least partlydisposed in external reactor 202. In a preferred embodiment, bothexternal and internal reactors 202 and 204 are bubble column reactors.Preferably, external reactor 202 includes an external reaction vessel206 and an external oxidant sparger 208, while internal reactor 204includes an internal reaction vessel 210 and an internal oxidant sparger212.

Although FIGS. 12 and 13 illustrate internal reaction vessel 210 asbeing fully disposed in external reaction vessel 206, it is possible forinternal reaction vessel 210 to be only partial disposed in externalreaction vessel 206. However, it is preferred for at least about 50, 90,95, or 100 percent of the height of internal reaction vessel 210 to belocated in external reaction vessel 206. Furthermore, it is preferredthat a portion of each reaction vessel is elevated above a portion ofthe other reaction vessel by at least about 0.01, 0.2, 1, or 2 times themaximum diameter of the external reaction vessel.

In a preferred embodiment of the present invention, external andinternal reaction vessels 206 and 210 each include a respective uprightsidewall having a generally cylindrical configuration. Preferably, theupright sidewalls of external and internal reaction vessels 206 and 210are substantially concentric and define an annulus therebetween.Internal reaction vessel 210 is supported vertically from externalreaction vessel 206, preferably principally by upright supports betweenthe lower portions of the respective vessels. In addition, internalreaction vessel 210 can be supported by external reaction vessel 206 viaa plurality of lateral support members 214 extending between the uprightsidewall of external and internal reaction vessels 206 and 210.Preferably, such lateral support members 214 have a non-foulingconfiguration with minimal upwardly-facing planar surface, as previouslydefined.

Although it is preferred for the upright sidewall of internal reactionvessel 210 to be substantially cylindrical, it is possible for certainportions of the upright sidewall of internal reaction vessel 210 to beconcave with respect to an adjacent portion of second reaction zone 218.Preferably, any portion of the upright sidewall of internal reactionvessel 210 that is concave with respect to an adjacent portion of secondreaction zone 218 accounts for less than about 25, 10, 5, or 0.1 percentof the total surface area of the upright sidewall of internal reactionvessel 210. Preferably, the ratio of the maximum height of the uprightsidewall of internal reaction vessel 210 to the maximum height of theupright sidewall of external reaction vessel 206 is in the range of fromabout 0.1:1 to about 0.9:1, more preferably in the range of from about0.2:1 to about 0.8:1, and most preferably in the range of from 0.3:1 to0.7:1.

External reaction vessel 206 defines therein a first reaction zone 216,while internal reaction vessel 210 defines therein a second reactionzone 218. Preferably, external and internal reaction vessels 206 and 210are aligned vertically such that the volumetric centroid of secondreaction zone 218 is horizontally displaced from the volumetric centroidof first reaction zone 216 by less than about 0.4, 0.2, 0.1, or 0.01times the maximum horizontal diameter of first reaction zone 216.Preferably, the ratio of the maximum horizontal cross sectional area offirst reaction zone 216 to second reaction zone 218 is in the range offrom about 0.01:1 to about 0.75:1, more preferably in the range of fromabout 0.03:1 to about 0.5:1, and most preferably in the range of from0.05:1 to 0.3:1. Preferably, the ratio of the horizontal cross sectionalarea of second reaction zone 218 to the horizontal cross sectional areaof the annulus defined between external and internal reaction vessels206 and 210 is at least about 0.02:1, more preferably in the range offrom about 0.05:1 to about 2:1, and most preferably in the range of fromabout 0.1:1 to about 1:1, where the cross sectional area is measured at¼-height, ½-height, and/or ¾-height of second reaction zone 218.Preferably, at least about 50, 70, 90, or 100 percent of the volume ofsecond reaction zone 218 is located in external reaction vessel 206.Preferably, the ratio of the volume of first reaction zone 216 to thevolume of second reaction zone 218 is in the range of from about 1:1 toabout 100:1, more preferably in the range of from about 4:1 to about50:1, and most preferably in the range of from 8:1 to 30:1. Preferably,first reaction zone 216 has a ratio of maximum vertical height tomaximum horizontal diameter in the range of from about 3:1 to about30:1, more preferably about 6:1 to about 20:1, and most preferably inthe range of from 9:1 to 15:1. Preferably, second reaction zone 218 hasa ratio of maximum vertical height to maximum horizontal diameter in therange of from about 0.3:1 to about 100:1, more preferably in the rangeof from about 1:1 to about 50:1, and most preferably in the range offrom 3:1 to 30:1. Preferably, the maximum horizontal diameter of secondreaction zone 218 is in the range of from about 0.1 to about 5 meters,more preferably in the range of from about 0.3 to about 4 meters, andmost preferably in the range of from 1 to 3 meters. Preferably, themaximum vertical height of second reaction zone 218 is in the range offrom about 1 to about 100 meters, more preferably in the range of fromabout 3 to about 50 meters, and most preferably in the range of from 10to 30 meters. Preferably, the ratio of the maximum horizontal diameterof second reaction zone 218 to the maximum horizontal diameter of firstreaction zone 216 is in the range of from about 0.05:1 to about 0.8:1,more preferably in the range of from about 0.1:1 to about 0.6:1, andmost preferably in the range of from 0.2:1 to 0.5:1. Preferably, theratio of the maximum vertical height of second reaction zone 218 to themaximum vertical height of first reaction zone 216 is in the range offrom about 0.03:1 to about 1:1, more preferably in the range of fromabout 0.1:1 to about 0.9:1, and most preferably in the range of from0.3:1 to 0.8:1. Any parameters (e.g., height, width, area, volume,relative horizontal placement, and relative vertical placement)specified herein for external reaction vessel 206 and appurtenances arealso construed as applying to first reaction zone 216 defined byexternal reaction vessel 206, and vice versa. Further, any parametersspecified herein for internal reaction vessel 210 and appurtenances arealso construed as applying to second reaction zone 218 defined byinternal reaction vessel 210, and vice versa.

During operation of bubble column reactor 200, a multi-phase reactionmedium 220 is first subjected to oxidation in first reaction zone 216and then subjected to oxidation in second reaction zone 218. Thus,during normal operation, a first portion of reaction medium 220 a islocated in first reaction zone 216, while a second portion of reactionmedium 220 b is located in second reaction zone 218. After beingprocessed in second reaction zone 218, a slurry phase (i.e., liquid andsolid phases) of reaction medium 220 b is withdrawn from second reactionzone 218 and discharged from bubble column reactor 200 via a slurryoutlet 222 for subsequent downstream processing.

Internal reactor 204 preferably comprises at least one internal gasopening that permits additional molecular oxygen to be discharged intosecond reaction zone 218. Preferably, a plurality of internal gasopenings are defined by internal oxidant sparger 212. The disclosuresfor oxidant sparger 34 of FIGS. 1-5 also apply to internal oxidantsparger 212 for conduit sizes and configurations, opening sizing andconfiguration, operating pressure drop, and liquid flushing. In notabledistinction, it is preferable to locate oxidant sparger 212 relativelyhigher in order to use a lower portion of internal reaction vessel 210as a deaeration zone. For example, embodiments disclosed herein foroxidation of para-xylene to form TPA provide a greatly diminished spacetime reaction rate near the bottom of second reaction zone 218, and thismitigates the effects of deaeration on impurity formation. Internalreaction vessel 210 has a maximum height “H_(i)”. It is preferred for atleast about 50, 75, 95, or 100 percent of the total open area defined byall of the internal gas openings to be spaced at least 0.05 H_(i), 0.1H_(i), or 0.25 H_(i) from the top of internal reaction vessel 210. It isalso preferred for at least about 50, 75, 95, or 100 percent of thetotal open area defined by all of the internal gas openings to be spacedless than about 0.5 H_(i), 0.25 H_(i), or 0.1 H_(i) above the bottom ofinternal reaction vessel 210. Preferably, at least about 50, 75, 95, or100 percent of the total open area defined by all of the internal gasopenings are spaced at least about 1, 5, or 10 meters from the top ofinternal reaction vessel 210 and at least about 0.5, 1, or 2 meters fromthe bottom of internal reaction vessel 210. It is preferred for at leastabout 50, 75, 95, or 100 percent of the total open area defined by allof the internal gas openings to communicate directly with secondreaction zone 218 and not communicate directly with first reaction zone216. As used herein, the term “open area” denotes the minimum surfacearea (planar or nonplanar) that would close off an opening.

In general, the manner in which the feed, oxidant, and reflux streamsare introduced into external reactor 202 and the manner in whichexternal reactor 202 is operated are substantially the same as describedabove with reference to bubble column reactor 20 of FIGS. 1-11. However,one difference between external reactor 202 (FIGS. 12 and 13) and bubblecolumn reactor 20 (FIGS. 1-11) is that external reactor 202 does notinclude an outlet that permits the slurry phase of reaction medium 220 ato be directly discharged from external reaction vessel 206 fordownstream processing. Rather, bubble column reactor 200 requires theslurry phase of reaction medium 220 a to first pass through internalreactor 204 before being discharged from bubble column reactor 200. Asmentioned above, in second reaction zone 218 of internal reactor 204,reaction medium 220 b is subjected to further oxidation to help purifythe liquid and/or solid phases of reaction medium 220 b.

In a process wherein para-xylene is fed to reaction zone 216, the liquidphase of reaction medium 220 a that exits first reaction zone 216 andenters second reaction zone 218 typically contains at least somepara-toluic acid. It is preferred for a substantial portion of thepara-toluic acid entering second reaction zone 218 to be oxidized insecond reaction zone 218. Thus, it is preferred for the time-averagedconcentration of para-toluic acid in the liquid phase of reaction medium220 b exiting second reaction zone 218 to be less than the time-averagedconcentration of para-toluic acid in the liquid phase of reaction medium220 a/b entering second reaction zone 218. Preferably, the time-averagedconcentration of para-toluic acid in the liquid phase of reaction medium220 b exiting second reaction zone 218 is less than about 50, 10, or 5percent of the time-averaged concentration of para-toluic acid in theliquid phase of reaction medium 220 a/b entering second reaction zone218. Preferably, the time-averaged concentration of para-toluic acid inthe liquid phase of reaction medium 220 a/b entering second reactionzone 218 is at least about 250 ppmw, more preferably in the range offrom about 500 to about 6,000 ppmw, and most preferably in the range offrom 1,000 to 4,000 ppmw. Preferably, the time-averaged concentration ofpara-toluic acid in the liquid phase of reaction medium 220 b exitingsecond reaction zone 218 is less than about 1,000, 250, or 50 ppmw.

Internal reaction vessel 210 is equipped with at least one directopening that permits reaction medium 220 a/b to pass directly betweenreaction zone 216 and second reaction zone 218. It is preferred forsubstantially all of the direct openings in internal reaction vessel 210to be located near the top of internal reaction vessel 210. Preferably,at least about 50, 75, 90, or 100 percent of the total open area definedby all of the direct openings is spaced less than about 0.5 H_(i), 0.25H_(i) or 0.1 H_(i) from the top of internal reaction vessel 210.Preferably, less than about 50, 25, 10, or 1 percent of the total openarea defined by the direct openings in internal reaction vessel 210 isspaced more than about 0.5 H_(i), 0.25 H_(i) or 0.1 H_(i) from the topof internal reaction vessel 210. Most preferably, the direct openingdefined by internal reaction vessel 210 is a single upper opening 224located at the upper-most end of internal reaction vessel 210. The ratioof the open area of upper opening 224 to the maximum horizontal crosssectional area of second reaction zone 218 is preferably at least about0.1:1, 0.2:1, or 0.5:1.

During normal operation of bubble column reactor 200, reaction medium220 passes from first reaction zone 216, through the direct opening(s)(e.g., upper opening 224) in internal reaction vessel 210, and intosecond reaction zone 218. In second reaction zone 218, the slurry phaseof reaction medium 220 b travels in a generally downward directionthrough second reaction zone 218, while the gas phase of reaction medium220 b travels in a generally upward direction. Preferably, internalreaction vessel 210 defines at least one discharge opening that permitsthe slurry phase to exit second reaction zone 218. The slurry phaseexiting the discharge opening of internal reaction vessel 210 then exitsbubble column reactor 200 via slurry outlet 222. Preferably, dischargeopening is located at or near the bottom of internal reaction vessel210. Preferably at least about 50, 75, 90, or 100 percent of the totalopen area defined by all discharge openings in internal reaction vessel210 is located within about 0.5 H_(i), 0.25 H_(i) or 0.1 H_(i) of thebottom of internal reaction vessel 210.

As reaction medium 220 b is processed in second reaction zone 218 ofinternal reactor 204, it is preferred for the gas hold-up of reactionmedium 220 b to decrease as the slurry phase of reaction medium 220 bflows downwardly through second reaction zone 218. Preferably, the ratioof the time-averaged gas hold-up of reaction medium 220 a/b enteringsecond reaction zone 218 to reaction medium 220 b exiting secondreaction zone 218 is at least about 2:1, 10:1, or 25:1. Preferably, thetime-averaged gas hold-up of reaction medium 220 a/b entering secondreaction zone 218 is in the range of from about 0.4 to about 0.9, morepreferably in the range of from about 0.5 to about 0.8, and mostpreferably in the range of from 0.55 to 0.7. Preferably, thetime-averaged gas hold-up of reaction medium 220 b exiting secondreaction zone 218 is less than about 0.1, 0.05, or 0.02. Preferably, theratio of the time-averaged gas hold-up of reaction medium 220 a in firstreaction zone 216 to reaction medium 220 b in second reaction zone 218is greater than about 1:1, more preferably in the range of from about1.25:1 to about 5:1, and most preferably in the range of from 1.5:1 to4:1, where the gas hold-up values are measured at any height of firstand second reaction zones 216 and 218, at any corresponding heights offirst and second reaction zones 216 and 218, at ¼-height of first and/orsecond reaction zones 216 and 218, at ½-height of first and/or secondreaction zones 216 and 218, at ¾-height of first and/or second reactionzones 216 and 218, and/or are average values over the entire heights offirst and/or second reaction zones 216 and 218. Preferably, thetime-averaged gas hold-up of the portion of reaction medium 220 a infirst reaction zone 216 is in the range of from about 0.4 to about 0.9,more preferably in the range of from about 0.5 to about 0.8, and mostpreferably in the range of from 0.55 to 0.70, where the gas hold-up ismeasured at any height of first reaction zone 216, at ¼-height of firstreaction zone 216, at ½-height of first reaction zone 216, at ¾-heightof first reaction zone 216, and/or is an average over the entire heightof first reaction zone 216. Preferably, the time-averaged gas hold-up ofthe portion of reaction medium 220 b in second reaction zone 218 is inthe range of from about 0.01 to about 0.6, more preferably in the rangeof from about 0.03 to about 0.3, and most preferably in the range offrom 0.08 to 0.2, where the gas hold-up is measured at any height ofsecond reaction zone 218, at ¼-height of second reaction zone 218, and½-height of second reaction zone 218, at ¾-height of second reactionzone 218, and/or is an average over the entire height of second reactionzone 218.

The temperature of reaction medium 220 is preferably approximately thesame in first and second reaction zones 216 and 218. Preferably, suchtemperature is in the range of from about 125 to about 200° C., morepreferably in the range of from about 140 to about 180° C., and mostpreferably in the range of from 150 to 170° C. However, temperaturedifferences preferably are formed within first reaction zone 216 thatare the same as disclosed herein with reference to FIG. 28. Preferably,temperature differences of the same magnitudes also exist within secondreaction zone 218 and also between first reaction zone 216 and secondreaction zone 218. These additional temperature gradients relate tochemical reaction occurring in second reaction zone 218, theintroduction additional oxidant to second reaction zone 218, and thestatic pressures extant in second reaction zone 218 compared to those infirst reaction zone 216. As disclosed above, the bubble hold-up ispreferably greater in first reaction zone 216 than in second reactionzone 218. Thus, at elevations below upper opening 224, the staticpressure in reaction zone 216 is greater than in second reaction zone218. The magnitude of this pressure difference depends on the magnitudeof liquid or slurry density and on the difference in bubble hold-upbetween the two reaction zones. The magnitude of this pressuredifference increases at elevations further below upper opening 224.

In one embodiment of the present invention, a portion of the oxidizablecompound (e.g., para-xylene) fed to bubble column reactor 200 isintroduced directly into second reaction zone 218 of internal reactor204. However, it is preferred for at least about 90, 95, 99, or 100 molepercent of the total oxidizable compound fed to bubble column reactor200 to be introduced into first reaction zone 216 (rather than secondreaction zone 218). Preferably, the molar ratio of the amount ofoxidizable compound introduced into first reaction zone 216 to theamount of oxidizable compound introduced into second reaction zone 218is at least about 2:1, 4:1, or 8:1.

Although FIGS. 12 and 13 depict a configuration where a portion of thetotal molecular oxygen fed to bubble column reactor 200 is introducedinto second reaction zone 218 of internal reactor 204 via internaloxidant sparger 212, it is preferred for the majority of the totalmolecular oxygen fed to bubble column reactor 200 to be introduced intofirst reaction zone 216, with the balance being introduced into thesecond reaction zone 218. Preferably, at least about 70, 90, 95, or 98mole percent of the total molecular oxygen fed to bubble column reactor200 is introduced into first reaction zone 216. Preferably, the molarratio of the amount of molecular oxygen introduced into first reactionzone 216 to the amount of molecular oxygen introduced into secondreaction zone 218 is at least about 2:1, more preferably in the range offrom about 4:1 to about 200:1, most preferably in the range of from 10:1to 100:1. Although it is possible for some of the solvent and/oroxidizable compound (e.g., para-xylene) to be fed directly to secondreaction zone 218, it is preferred for less than about 10, 5, or 1weight percent of the total amount of solvent and/or oxidizable compoundfed to bubble column reactor 200 to be fed directly to second reactionzone 218.

The volume, residence time, and space time rate of medium 220 a in firstreaction zone 216 of external reaction vessel 206 are preferablysubstantially greater than the volume, residence time, and space timerate of reaction medium 220 b in second reaction zone 218 of internalreaction vessel 210. Therefore, the majority of the oxidizable compound(e.g., para-xylene) fed to bubble column reactor 200 is preferablyoxidized in first reaction zone 216. Preferably, at least about 80, 90,or 95 weight percent of all the oxidizable compound that is oxidized inbubble column reactor 200 is oxidized in first reaction zone 216. It ispreferred for the time-averaged superficial gas velocity of reactionmedium 220 a in first reaction zone 216 to be at least about 0.2, 0.4,0.8, or 1 meters per second, where the superficial gas velocity ismeasured at any height of first reaction zone 216, at ¼-height of firstreaction zone 216, at ½-height of first reaction zone 216, at ¾-heightof first reaction zone 216, and/or is an average over the entire heightof first reaction zone 216.

Although reaction medium 220 b in second reaction zone 218 can have thesame superficial gas velocity as reaction medium 220 a in first reactionzone 216, it is preferred that the time-averaged superficial gasvelocity of reaction medium 220 b in second reaction zone 218 is lessthan the time-averaged and volume-averaged superficial gas velocity ofreaction medium 220 b in second reaction zone 218. This reducedsuperficial gas velocity in second reaction zone 218 is made possibleby, for example, the reduced demand for molecular oxygen in secondreaction zone 218 compared to first reaction zone 216. Preferably, theratio of the time-averaged superficial gas velocity of reaction medium220 a in first reaction zone 216 to reaction medium 220 b in secondreaction zone 218 is at least about 1.25:1, 2:1, or 5:1, where thesuperficial gas velocities are measured at any height of first andsecond reaction zones 216 and 218, at any corresponding heights of firstand second reaction zones 216 and 218, at ¼-height of first and/orsecond reaction zones 216 and 218, at ½-height of first and/or secondreaction zones 216 and 218, at ¾-height of first and/or second reactionzones 216 and 218, and/or are average values over the entire heights offirst and/or second reaction zones 216 and 218. Preferably, thetime-averaged and volume-averaged superficial gas velocity of reactionmedium 220 b in second reaction zone 218 is less than about 0.2, 0.1, or0.06 meters per second, where the superficial gas velocity is measuredat any height of second reaction zone 218, at ¼-height of secondreaction zone 218, at ½-height of second reaction zone 218, at ¾-heightof second reaction zone 218, and/or is an average over the entire heightof second reaction zone 218. With these lower superficial gasvelocities, downward flow of the slurry phase of reaction medium 220 bin second reaction zone 218 can be made to move directionally towardplug flow. For example, during oxidation of para-xylene to form TPA, therelative vertical gradient of liquid phase concentration of para-toluicacid can be much greater in second reaction zone 218 than in firstreaction zone 216. This is notwithstanding that second reaction zone 218is a bubble column having axial mixing of liquid and of slurrycompositions. The time-averaged superficial velocity of the slurry phase(solid+liquid) and the liquid phase of reaction medium 220 b in secondreaction zone 218 are preferably less than about 0.2, 0.1, or 0.06meters per second, where the superficial velocity is measured at anyheight of second reaction zone 218, at ¼-height of second reaction zone218, at ½-height of second reaction zone 218, at ¾-height of secondreaction zone 218, and/or is an average over the entire height of secondreaction zone 218.

In one embodiment of the present invention, bubble column reactor 200 isoperated in a manner that permits solids sedimentation in internalreactor 204. If solids sedimentation is desired, it is preferred for thetime-averaged and volume-averaged superficial gas velocity of reactionmedium 220 b in second reaction zone 218 to be less than about 0.05,0.03, or 0.01 meters per second.

Further, if solids sedimentation is desired, it is preferred for thetime-averaged and volume-averaged superficial velocity of the slurry andliquid phases of reaction medium 220 b in second reaction zone 218 to beless than about 0.01, 0.005, or 0.001 meters per second.

While it is possible for some of the slurry phase exiting internalreactor 204 to be directly recirculated back to first reaction zone 216without further downstream processing, it is preferred for directrecirculation of reaction medium 220 b from the lower elevations ofsecond reaction zone 218 to first reaction zone 216 to be minimizedPreferably, the mass of reaction medium 220 b (solid, liquid, and gasphases) exiting the lower 25 percent of the volume of second reactionzone 218 and directly recirculated back to first reaction zone 216without further downstream processing is less than 10, 1, or 0.1 timesthe mass (solid, liquid, and gas phases) of reaction medium 220 bexiting second reaction zone 218 and thereafter subjected to downstreamprocessing. Preferably, the mass of reaction medium 220 b exiting thelower 50 percent of the volume of second reaction zone 218 and directlyrecirculated back to first reaction zone 216 without further downstreamprocessing is less than 20, 2, or 0.2 times the mass of reaction medium220 b exiting second reaction zone 218 and thereafter subjected todownstream processing. Preferably, less than about 50, 75, or 90 weightpercent of the liquid phase of reaction medium 220 b exiting secondreaction zone 218 via openings in the lower 90, 60, 50, or 5 percent ofthe volume of second reaction zone 218 is introduced into first reactionzone 216 within 60, 20, 5, or 1 minutes after exiting second reactionzone 218. Preferably, the liquid phase of reaction medium 220 b locatedin second reaction zone 218 has a mass-averaged residence time in secondreaction zone 218 of at least about 1 minute, more preferably in therange of from about 2 to about 60 minutes, and most preferably in therange of from 5 to 30 minutes. Preferably, less than about 50, 75, or 90weight percent of the liquid phase of reaction medium 220 a/b introducedinto second reaction zone 218 enters second reaction zone 218 in thelower 90, 60, or 30 percent of the volume of second reaction zone 218.Preferably, less than about 50, 75, or 90 weight percent of the totalliquid phase of reaction medium 220 a/b introduced as a liquid-phasefeed stream into first reaction zone 216 enters first reaction zone 216within 60, 20, 5, or 1 minutes after being withdrawn from secondreaction zone 218 via slurry outlet 222. Preferably, at least about 75,90, 95, or 99 weight percent of the liquid phase of reaction medium 220b withdrawn from second reaction zone 218 exits second reaction zone 218via openings in the lower 90, 60, 30, or 5 percent of the volume ofsecond reaction zone 218.

The design of reactor-in-reactor bubble column reactor 200 can be variedin many ways without departing from the ambit of the present invention.For example, internal reaction vessel 210 can have a greater height thanexternal reaction vessel 206 if internal reaction vessel 210 extendsbelow the lower end of external reaction vessel 206. External andinternal reaction vessels 206 and 210 can be cylindrical, asillustrated, or can have another shape. External and internal reactionvessels 206 and 210 need not be axisymmetric, axially vertical, orconcentric. The gas phase exiting internal reactor 204 can be routedoutside bubble column reactor 200 without being commingled with reactionmedium 220 a in first reaction zone 216. However, for flammabilitysafety, it is desirable to limit volumes of trapped gas pockets to lessthan about 10, 2, or 1 cubic meters. In addition, the slurry phaseexiting internal reactor 204 need not exit via a single slurry openingin the bottom of internal reaction vessel 210. The slurry phase can exitbubble column reactor 200 though a side outlet in a pressure containingsidewall of external reactor 202.

Referring now to FIG. 14, there is illustrated a bubble column reactor300 having a reactor-in-reactor and staged-diameter configuration.Bubble column reactor 300 comprises an external reactor 302 and aninternal reactor 304. External reactor 302 includes an external reactionvessel 306 having a broad lower section 306 a and a narrow upper section306 b. Preferably, the diameter of narrow upper section 306 b is smallerthan the diameter of broad lower section 306 a. With the exception ofthe staged-diameter configuration of the external reaction vessel,bubble column reactor 300 of FIG. 14 is preferably configured andoperated in substantially the same manner as bubble column reactor 200of FIGS. 12 and 13, described above.

Referring now to FIG. 15, there is illustrated a reactor system 400comprising a primary oxidation reactor 402 and a secondary oxidationreactor 404. Primary oxidation reactor 402 is preferably configured andoperated in substantially the same manner as external reactor 202 ofFIGS. 12 and 13.

Secondary oxidation reactor 404 is preferably configured and operated insubstantially the same manner as internal reactor 204 of FIGS. 12 and13. However, the main difference between reactor system 400 of FIG. 15and bubble column reactor 200 of FIGS. 12 and 13 is that secondaryoxidation reactor 404 of reactor system 400 is located outside ofprimary oxidation reactor 402. In reaction system 400 of FIG. 15, aninlet conduit 405 is employed to transfer a portion of the reactionmedium 420 from primary oxidation reactor 402 to secondary oxidationreactor 404. Further, an outlet conduit 407 is used to transfer overheadgasses from the top of secondary oxidation reactor 404 to primaryoxidation reactor 402.

During normal operation of reaction system 400, reaction medium 420first undergoes oxidation in a primary reaction zone 416 of primaryoxidation reactor 402. Reaction medium 420 a is then withdrawn fromprimary reaction zone 416 and transferred to a secondary reaction zone418 via conduit 405. In secondary reaction zone 418, the liquid and/orsolid phases of reaction medium 420 b are subjected to furtheroxidation. It is preferred for at least about 50, 75, 95, or 99 weightpercent of liquid and/or solid phases withdrawn from primary reactionzone 416 to be processed in secondary reaction zone 416. Overhead gassesexit an upper gas outlet of secondary oxidation reactor 404 and aretransferred back to primary oxidation reactor 402 via conduit 407. Aslurry phase of reaction medium 420 b exits a lower slurry outlet 422 ofsecondary oxidation reactor 404 and is thereafter subjected to furtherdownstream processing.

Inlet conduit 405 may attach to primary oxidation reactor 402 at anyheight. Although not shown in FIG. 15, reaction medium 420 can bemechanically pumped to secondary reaction zone 418 if desired. However,it is more preferable to use elevation head (gravity) to transferreaction medium 420 from primary reaction zone 416 through inlet conduit405 and into secondary reaction zone 418. Accordingly it is preferablethat inlet conduit 405 is connected on one end to the upper 50, 30, 20,or 10 percent of the total height and/or volume of primary reaction zone416. Preferably, the other end of inlet conduit 405 is attached to theupper 30, 20, 10, or 5 percent of the total height and/or volume ofsecondary reaction zone 418. Preferably, inlet conduit 405 is horizontaland/or sloping downward from primary oxidation reactor 402 towardsecondary oxidation reactor 404. Outlet conduit 407 may attach to anyelevation in secondary oxidation reactor 404, but it is preferable thatoutlet conduit 407 is connected to secondary oxidation reactor 404 abovethe attachment elevation of inlet conduit 405. More preferably, outletconduit 407 attaches to the top of secondary oxidation reactor 404.Outlet conduit 407 preferably attaches to primary oxidation reactor 402above the attachment elevation of inlet conduit 405. More preferably,outlet conduit 407 attaches to the upper 30, 20, 10, or 5 percent of thetotal height and/or volume of primary reaction zone 416. Preferably,outlet conduit 407 is horizontal and/or sloping upward from reactionsecondary oxidation reactor 404 toward primary oxidation reactor 402.Although not shown in FIG. 15, outlet conduit 407 may also attachdirectly to the gas outlet conduit that withdraws gaseous effluent fromthe top of primary oxidation reactor 402. The upper extent of secondaryreaction zone 416 may be above or below the upper extent of primaryreaction zone 418. More preferably, the upper extent of primary reactionzone 416 is within 10 meters above to 50 meters below, 2 meters below to40 meters below, or 5 meters below to 30 meters below the upper extentof secondary reaction zone 418. Lower slurry outlet 422 may exit fromany elevation of secondary oxidation reactor 404, but it is preferablethat lower slurry outlet 422 is connected to secondary oxidation reactor404 below the attachment elevation of inlet conduit 405. The attachmentpoint of lower slurry outlet 422 is more preferably widely separated inelevation from the attachment point of inlet conduit 405, with the twoattachments separated by at least about 50, 70, 90, or 95 percent of theheight of secondary reaction zone 418. Most preferably, lower slurryoutlet 422 attaches to the bottom of secondary oxidation reactor 404 asshown in FIG. 15. The lower extent of secondary reaction zone 418 may beelevated above or below the lower extent of primary reaction zone 416.More preferably, the lower extent of primary reaction zone 416 iselevated within about 40, 20, 5, or 2 meters above or below the lowerextent of secondary reaction zone 418.

Parameters (e.g., height, width, area, volume, relative horizontalplacement, and relative vertical placement) specified herein for primaryoxidation reactor 402 and appurtenances are also construed as applyingto primary reaction zone 416 defined by primary oxidation reactor 402,and vice versa. Any parameters specified herein for secondary oxidationreactor 404 and appurtenances are also construed as applying tosecondary reaction zone 418 defined by secondary oxidation reactor 404,and vice versa.

As mentioned above, it is preferred for secondary oxidation reactor 404to be located outside of primary oxidation reactor 402. Preferably,secondary oxidation reactor 404 is located alongside primary oxidationreactor 402 (i.e., at least a portion of primary and secondary oxidationreactors 402 and 404 share a common elevation). Primary reaction zone416 of primary oxidation reactor 402 has a maximum diameter “D_(p)”. Thevolumetric centroid of secondary reaction zone 418 is preferablyhorizontally spaced from the volumetric centroid of primary reactionzone 416 by at least about 0.5 D_(p), 0.75 D_(p), or 1.0 D_(p) and byless than about 30 D_(p), 10 D_(p), or 3 D_(p).

Referring now to FIG. 16, there is illustrated a reactor system 500comprising a primary oxidation reactor 502 and a secondary oxidationreactor 504. Primary oxidation reactor defines therein a primaryoxidation zone 516, while secondary oxidation reactor 504 definestherein a secondary oxidation zone 518. Each reaction zone 516 and 518receives a portion of reaction medium 520.

The configuration and operation of reactor system 500 (FIG. 16) ispreferably substantially the same as the configuration and of reactorsystem 400 (FIG. 15). However, in reactor system 500, the uprightsidewall of primary oxidation reactor 502 defines at least one enlargedopening 505 that permits the transfer of reaction medium 520 fromprimary reaction zone 516 to secondary reaction zone 518, whilesimultaneously permitting the transfer of the disengaged gas phase fromsecondary reaction zone 518 to primary reaction zone 516. Preferably,the open area of enlarged opening 505 divided by the maximum horizontalcross sectional area of the upright portion of secondary reaction zone218 is in the range of from about 0.01 to 2, 0.02 to 0.5, or 0.04 to0.2. Primary reaction zone 516 of primary oxidation reactor 502 has amaximum height “H_(p)”. It is preferred for the areal center of enlargedopening 505 to be vertically spaced at least about 0.1 H_(p), 0.2 H_(p),or 0.3 H_(p) from the top and/or bottom of primary reaction zone 516.

Referring now to FIGS. 17-25, there is illustrated a number of bubblecolumn reactors equipped with internal structures having a variety ofconfigurations. It has been discovered that employing one or moreinternal structures surrounded by the reaction medium surprisinglymodifies end-to-end mixing of the reaction medium. The internalstructure defines a quiescent zone having reduced turbulence compared tothe turbulence of the reaction medium surrounding the quiescent zone.

As illustrated in FIGS. 17-25, the internal structure can take a varietyof forms. In particular, FIG. 17 illustrates a bubble column reactor 600that employs a generally cylindrical internal structure 602 to definethe quiescent zone. Internal structure 602 is substantially centered inthe main reaction zone of bubble column reactor 600 and is verticallyspaced from the top and bottom ends of the main reaction zone. FIG. 18illustrates a bubble column reactor 610 that employs a generallycylindrical internal structure 612 that is similar to internal structure602 of FIG. 17. However, internal structure 612 of FIG. 18 is notcentered in the main reaction zone of bubble column reactor 610. Rather,the volumetric centroid of the quiescent zone defined by internalstructure 612 is horizontally offset from the volumetric centroid of themain reaction zone.

Further, the bottom of internal structure 612 is located near the lowertangent line of bubble column reactor 610. FIG. 19 illustrates a bubblecolumn reactor 620 employing a generally cylindrical internal structure622 that is taller than the internal structure 602 and 612 of FIGS. 17and 18. Further, the volumetric centroid of the quiescent zone definedby internal structure 622 is offset from the volumetric centroid of themain reaction zone of bubble column reactor 620. FIG. 20 illustrates abubble column reactor 630 employing an internal structure comprising agenerally cylindrical upper portion 632 and a generally cylindricallower portion 634. Lower portion 634 of the internal structure has anarrower diameter than upper portion 632. FIG. 21 illustrates a bubblecolumn reactor 640 employing an internal structure comprising agenerally cylindrical lower portion 642 and a generally cylindricalupper portion 644. Upper portion 644 of the internal structure has anarrower diameter than lower portion 642. FIG. 22 illustrates a bubblecolumn reactor 650 employing first, second, and third separate internalstructures 652, 654, and 656. Internal structures 652, 654, and 656 arevertically spaced from one another. The volumetric centroids of thequiescent zones defined by first and third internal structures 652 and656 are horizontally aligned with the volumetric centroid of the mainreaction zone of bubble column reactor 650. However, the volumetriccentroid of the quiescent zone defined by second internal structure 654is horizontally offset from the volumetric centroid of the main reactionzone of bubble column reactor 650. FIG. 23 illustrates a bubble columnreactor 660 employing a pair of side-by-side first and second internalstructures 662 and 664. The volumetric centroids of the quiescent zonesdefined by first and second internal structures 662 and 664 arehorizontally spaced from one another and horizontally spaced from thevolumetric centroid of the main reaction zone of bubble column reactor660. Further, first and second internal structures 662 and 664 have aside-by-side configuration so that at least a portion of first andsecond internal structures 662 and 664 share a common elevation. FIG. 24illustrates a bubble column reactor 760 employing a generally prismaticinternal structure 672. In particular, internal structure 672 has agenerally triangular horizontal cross section. FIG. 25 illustrates abubble column reactor 680 employing a generally cylindrical internalstructure 682 that is similar to internal structure 602 of FIG. 17.However, the external reaction vessel of bubble column reactor 680 has astepped diameter created by a narrow lower section 682 and a broad uppersection 684.

As illustrated in FIGS. 17-25, the internal structure employed inaccordance with one embodiment of the present invention can have avariety of shapes and can be disposed in a variety of positions withinthe main reaction zone of the bubble column reactor. Further, theinternal structure and the quiescent zone defined therein can be formedof a variety of different materials. In one embodiment of the presentinvention, the internal structure is completely closed, so that none ofthe surrounding reaction medium enters the internal structure. Such aclosed internal structure can be hollow or solid. In another embodimentof the present invention, the internal structure includes one or moreopenings that allow the reaction medium to enter the quiescent zonedefined by the internal structure. However, because one purpose of thequiescent zone is to create a zone of reduced turbulence relative to theturbulence of the reaction medium surrounding it, it is preferred thatthe internal structure does not allow a significant amount of thereaction medium to rapidly flow through the internal structure.

The specific configuration and operating parameters of a bubble columnreactor equipped with one or more internal structures will now bedescribed in greater detail. Preferably, the internal structure isdisposed entirely inside of the external reaction vessel of the bubblecolumn reactor; however, it is possible for at least a portion of theinternal structure to protrude outside of the external reaction vesselof the bubble column reactor. As mentioned above, during operation ofthe bubble column reactor, the internal structure defines at least onequiescent zone within the bubble column reactor. The main reaction zoneof the bubble column reactor and the quiescent zone are distinct volumes(i.e., do not overlap one another). The main reaction zone of the bubblecolumn reactor is defined inside the external reaction vessel of thebubble column reactor, but outside of the internal structure.

As mentioned above, the quiescent zone defined by the internal structureis a volume that has reduced turbulence relative to the turbulence ofthe adjacent reaction medium in the main reaction zone. It is preferredfor at least about 90, 95, 98, or 99.9 percent of the volume of thequiescent zone to be filled with a material other than the reactionmedium and/or to be filled with a portion of the reaction medium havingsubstantially reduced turbulence compared to the reaction medium locatedadjacent the internal structure. If the quiescent zone includes anyportion of the reaction medium, it is preferred for the portion of thereaction medium contained in the quiescent zone to have a mass-averagedresidence time in the quiescent zone of at least about 2, 8, 30, or 120minutes. If the quiescent zone includes any portion of the reactionmedium, it is preferred for the time-averaged gas hold-up of thereaction medium in the quiescent zone to be less than about 0.2, 0.1,0.5, or 0.01, where the gas hold-up is measured at any elevation of thequiescent zone, ¼-height of the quiescent zone, ½-height of thequiescent zone, ¾-height of the quiescent zone, and/or is an averageover the entire height of the quiescent zone. It is preferred for thetime-averaged gas hold-up of the reaction medium in the reaction zone tobe in the range of from about 0.2 to about 0.9, more preferably, about0.5 to about 0.8, and most preferably, 0.55 to 0.7, where the gashold-up is measured at any elevation of the reaction zone, ¼-height ofthe reaction zone, ½-height of the reaction zone, ¾-height of thereaction zone, and/or is an average over the entire height of thereaction zone. If the quiescent zone includes any portion of thereaction medium, it is preferred for the time-averaged superficial gasvelocity of the reaction medium in the quiescent zone to be less thanabout 0.4, 0.2, 0.1, or 0.05 meters per second, where the superficialgas velocity is measured at any elevation of the quiescent zone,¼-height of the quiescent zone, ½-height of the quiescent zone, ¾-heightof the quiescent zone, and/or is an average over the entire height ofthe quiescent zone. It is preferred for the time-averaged superficialgas velocity of the reaction medium in the reaction zone to be at leastabout 0.2, 0.4, 0.8, or 1 meters per second, where the superficial gasvelocity is measured at any elevation of the reaction zone, ¼-height ofthe reaction zone, ½-height of the reaction zone, ¾-height of thereaction zone, and/or is an average over the entire height of thereaction zone. If the quiescent zone includes any portion of thereaction medium, it is preferred for the time-averaged superficialvelocity of the liquid phase of the reaction medium in the quiescentzone to be less than about 0.04, 0.01, or 0.004 meters per second, wherethe superficial velocity of the liquid phase is measured at anyelevation of the quiescent zone, ¼-height of the quiescent zone,½-height of the quiescent zone, ¾-height of the quiescent zone, and/oris an average over the entire height of the quiescent zone. It ispreferred for the time-averaged superficial velocity of the liquid phaseof the reaction medium in the reaction zone to be less than about 0.1,0.04, or 0.01 meters per second, where the superficial velocity of theliquid phase is measured at any elevation of the reaction zone, ¼-heightof the reaction zone, ½-height of the reaction zone, ¾-height of thereaction zone, and/or is an average over the entire height of thereaction zone. Any parameters (e.g., height, width, area, volume,relative horizontal placement, and relative vertical placement)specified herein for the internal structure are also construed asapplying to the quiescent zone defined by the internal structure, andvice versa.

It is preferred for the size of the quiescent zone defined by theinternal structure to be such that the quiescent zone includes thereinat least one location that is spaced from the reaction zone by at leastabout 0.05 times the maximum horizontal diameter of the reaction zone orabout 0.2 meters, whichever is larger. Preferably, the quiescent zoneincludes therein at least one location that is spaced from the reactionzone by at least about 0.4, 0.7, or 1.0 meters. Preferably, thequiescent zone includes therein at least one location that is spacedfrom the reaction zone by at least about 0.1, 0.2, or 0.3 times themaximum horizontal diameter of the reaction zone. The quiescent zonepreferably includes therein at least two locations that are spaced fromone another by a vertical distance that is at least about 0.5, 1, 2, or4 times the maximum horizontal diameter of the reaction zone.Preferably, these two vertically-spaced locations in the quiescent zoneare also each separated from the reaction zone by at least about 0.05,0.1, 0.2, or 0.3 times the maximum horizontal diameter of the reactionzone. Preferably, these two vertically-spaced locations in the quiescentzone are vertically-spaced from one another by at least about 1, 3, 10,or 20 meters and are each also separated from the reaction zone by atleast about 0.1, 0.4, 0.7, or 1 meters. Preferably, the volume of thequiescent zone is in the range of from about 1 to about 50 percent ofthe volume of the main reaction zone, more preferably in the range offrom about 2 to about 25 percent of the volume of the main reactionzone, and most preferably in the range of from 4 to 15 percent of thevolume of the main reaction zone.

The external reaction vessel of the bubble column reactor preferablycomprises a generally cylindrical upright external sidewall. Preferably,the internal structure comprises a generally cylindrical uprightinternal sidewall that is spaced inwardly from the external sidewall.Preferably, the internal structure is not part of a heat exchanger.Thus, it is preferred for the time-averaged heat flux through theupright internal sidewalls of the internal structure to be less thanabout 100, 15, 3, or 0.3 kilowatts per square meter. An annulus filledwith the reaction medium is preferably defined between the internal andexternal sidewalls. The internal structure is supported vertically fromthe external vessel, preferably by upright supports between the lowerportions of internal structure and the lower portion of externalreaction vessel. In addition, the internal structure is preferablysupported by the external reaction vessel via a plurality of non-foulinglateral support members extending inwardly from the external sidewall tothe internal sidewall. Preferably, the horizontal cross sectional areaof the quiescent zone at ¼-height, ½-height, and/or ¾-height of thequiescent zone is at least about 2, 5 to 75, or 10 to 30 percent of thehorizontal cross sectional area of the annulus at the respectiveelevations. Preferably, the maximum height of the internal uprightsidewall is in the range of from about 10 to about 90 percent of themaximum height of the external upright sidewall, more preferably in therange of from about 20 to about 80 percent of the maximum height of theexternal upright sidewall, and most preferably in the range of 30 to 70percent of the maximum height of the external upright sidewall. Althoughit is preferred for the internal sidewall to have a generallycylindrical configuration, it is possible that a portion of the internalsidewall may be concave with respect to an adjacent portion of thequiescent zone. When the internal sidewall includes a concave portion,it is preferred for this concave portion to form less than about 25, 10,5, or 0.1 percent of the total outwardly facing surface area presentedby the internal sidewall. Preferably, the ratio of the total surfacearea of the internal structure that is in direct contact with thereaction medium to the total volume of the reaction zone is less thanabout 1, 0.5, 0.3, or 0.15 meters square per cubic meter. It ispreferred for the volumetric centroid of the quiescent zone to behorizontally displaced from the volumetric centroid of the main reactionzone by less than about 0.4, 0.2, 0.1, or 0.01 times the maximumhorizontal diameter of the main reaction zone.

When the bubble column reactor includes more than one internal structuredefining more than one quiescent zone, it is preferred for the quiescentzones to be vertically aligned such that the volumetric centroid of allthe quiescent zones considered together is horizontally displaced fromthe volumetric centroid of the reaction zone by less than about 0.4,0.2, 0.1, or 0.01 times the maximum horizontal diameter of the mainreaction zone. Further, when a plurality of quiescent zones are formedwithin the main reaction zone, it is preferred for the number ofindividual quiescent zones having a volume greater than 0.2 percent ofthe volume of the main reaction zone to be less than about 100, 10, 5,or 2.

The external reaction vessel of the bubble column reactor preferably hasa ratio of maximum vertical height to maximum horizontal diameter in therange of from about 3:1 to about 30:1, more preferably in the range offrom about 6:1 to about 20:1, and most preferably in the range of from9:1 to 15:1. The internal structure preferably has a ratio of maximumvertical height to maximum horizontal diameter in the range of fromabout 0.3:1 to about 100:1, more preferably in the range of from about1:1 to about 50:1, and most preferably in the range of from 3:1 to 30:1.It is preferred for the maximum horizontal diameter of the internalstructure to be in the range of from about 0.1 to about 5 meters, morepreferably in the range of from about 0.3 to about 4 meters, and mostpreferably in the range of from 1 to 3 meters. Preferably, the maximumvertical height of the internal structure is in the range of from about1 to about 100 meters, more preferably in the range of from about 3 toabout 50 meters, and most preferably in the range of from 10 to 50meters. Preferably, the maximum horizontal diameter of the internalstructure is in the range of from about 5 to about 80, more preferablyabout 10 to about 60, and most preferably 20 to 50 percent of themaximum horizontal diameter of the external reaction vessel. Preferably,the maximum vertical height of internal structure 602 is in the range offrom about 3 to about 100 percent of the maximum vertical height of theexternal reaction vessel, more preferably in the range of from about 10to about 90 percent of the maximum vertical height of the externalreaction vessel, and most preferably in the range of from 30 to 80percent of the maximum vertical height of the external reaction vessel.Any parameters (e.g., height, width, area, volume, relative horizontalplacement, and relative vertical placement) specified herein for theexternal reaction vessel and appurtenances are also construed asapplying to the reaction zone defined by the external reaction vessel,and vice versa.

In one embodiment of the present invention, the internal structurecompletely isolates the quiescent zone from the reaction zone. In analternative embodiment, the internal structure defines one or moredirect openings that permit direct fluid communication between thequiescent zone and the reaction zone. When the internal structuredefines such direct openings, it is preferred for the maximum diameterof the smallest of the direct openings to be less than about 0.3, 0.2,0.1, or 0.05 times the maximum horizontal diameter of the main reactionzone. When the internal structure defines such direct openings, it ispreferred for the maximum diameter of the largest of the direct openingsto be less than about 0.4, 0.3, 0.2, or 0.1 times the maximum horizontaldiameter of the main reaction zone. When the internal structure definessuch direct openings, it is preferred for the cumulative open areadefined by all of the direct openings to be less than about 0.4, 0.3, or0.2 times the maximum horizontal cross sectional area of the mainreaction zone. The internal structure has a maximum height (H_(i)). Whenthe internal structure defines one or more direct openings, it ispreferred that less than about 50, 25, or 10 percent of the cumulativeopen area defined by all of the direct openings is spaced more thanabout 0.5 H₁, 0.25 H_(i), or 0.1 H_(i) from the top of the internalstructure. When the bubble column reactor employs a plurality ofinternal structures to form a plurality of distinct quiescent zones, itis possible for two or more of the quiescent zones to includeinterconnecting openings and/or conduits that permit fluid communicationbetween quiescent zones. Preferably, the maximum diameter of thesmallest of each of these interconnected openings and/or conduits isless than about 0.3, 0.2, 0.1, or 0.05 times the maximum horizontaldiameter of the main reaction zone.

As mentioned above, certain physical and operational features of thebubble column reactors, described above with reference to FIGS. 1-25,provide for vertical gradients in the pressure, temperature, andreactant (i.e., oxygen and oxidizable compound) concentrations of theprocessed reaction medium. As discussed above, these vertical gradientscan provide for a more effective and economical oxidation process ascompared to conventional oxidations processes, which favor a well-mixedreaction medium of relatively uniform pressure, temperature, andreactant concentration throughout. The vertical gradients for oxygen,oxidizable compound (e.g., para-xylene), and temperature made possibleby employing an oxidation system in accordance with an embodiment of thepresent invention will now be discussed in greater detail.

Referring now to FIG. 26, in order to quantify the reactantconcentration gradients existing in the reaction medium during oxidationin the bubble column reactor, the entire volume of the reaction mediumcan be theoretically partitioned into 30 discrete horizontal slices ofequal volume. FIG. 26 illustrates the concept of dividing the reactionmedium into 30 discrete horizontal slices of equal volume. With theexception of the highest and lowest horizontal slices, each horizontalslice is a discrete volume bounded on its top and bottom by imaginaryhorizontal planes and bounded on its sides by the wall of the reactor.The highest horizontal slice is bounded on its bottom by an imaginaryhorizontal plane and on its top by the upper surface of the reactionmedium. The lowest horizontal slice is bounded on its top by animaginary horizontal plane and on its bottom by the bottom of the vesselshell. Once the reaction medium has been theoretically partitioned into30 discrete horizontal slices of equal volume, the time-averaged andvolume-averaged concentration of each horizontal slice can then bedetermined. The individual horizontal slice having the maximumconcentration of all 30 horizontal slices can be identified as the“C-max horizontal slice.” The individual horizontal slice located abovethe C-max horizontal slice and having the minimum concentration of allhorizontal slices located above the C-max horizontal slice can beidentified as the “C-min horizontal slice.” The vertical concentrationgradient can then be calculated as the ratio of the concentration in theC-max horizontal slice to the concentration in the C-min horizontalslice.

With respect to quantifying the oxygen concentration gradient, when thereaction medium is theoretically partitioned into 30 discrete horizontalslices of equal volume, an O₂-max horizontal slice is identified ashaving the maximum oxygen concentration of all the 30 horizontal slicesand an O₂-min horizontal slice is identified as having the minimumoxygen concentration of the horizontal slices located above the O₂-maxhorizontal slice. The oxygen concentrations of the horizontal slices aremeasured in the gas phase of the reaction medium on a time-averaged andvolume-averaged molar wet basis. It is preferred for the ratio of theoxygen concentration of the O₂-max horizontal slice to the oxygenconcentration of the O₂-min horizontal slice to be in the range of fromabout 2:1 to about 25:1, more preferably in the range of from about 3:1to about 15:1, and most preferably in the range of from 4:1 to 10:1.

Typically, the O₂-max horizontal slice will be located near the bottomof the reaction medium, while the O₂-min horizontal slice will belocated near the top of the reaction medium. Preferably, the O₂-minhorizontal slice is one of the 5 upper-most horizontal slices of the 30discrete horizontal slices. Most preferably, the O₂-min horizontal sliceis the upper-most one of the 30 discrete horizontal slices, asillustrated in FIG. 26. Preferably, the O₂-max horizontal slice is oneof the 10 lower-most horizontal slices of the 30 discrete horizontalslices. Most preferably, the O₂-max horizontal slice is one of the 5lower-most horizontal slices of the 30 discrete horizontal slices. Forexample, FIG. 26 illustrates the O₂-max horizontal slice as the thirdhorizontal slice from the bottom of the reactor. It is preferred for thevertical spacing between the O₂-min and O₂-max horizontal slices to beat least about 2 W, more preferably at least about 4 W, and mostpreferably at least 6 W. It is preferred for the vertical spacingbetween the O₂-min and O₂-max horizontal slices to be at least about 0.2H, more preferably at least about 0.4 H, and most preferably at least0.6 H

The time-averaged and volume-averaged oxygen concentration, on a wetbasis, of the O₂-min horizontal slice is preferably in the range of fromabout 0.1 to about 3 mole percent, more preferably in the range of fromabout 0.3 to about 2 mole percent, and most preferably in the range offrom 0.5 to 1.5 mole percent. The time-averaged and volume-averagedoxygen concentration of the O₂-max horizontal slice is preferably in therange of from about 4 to about 20 mole percent, more preferably in therange of from about 5 to about 15 mole percent, and most preferably inthe range of from 6 to 12 mole percent. The time-averaged concentrationof oxygen, on a dry basis, in the gaseous effluent discharged from thereactor via the gas outlet is preferably in the range of from about 0.5to about 9 mole percent, more preferably in the range of from about 1 toabout 7 mole percent, and most preferably in the range of from 1.5 to 5mole percent.

Because the oxygen concentration decays so markedly toward the top ofthe reaction medium, it is desirable that the demand for oxygen bereduced in the top of the reaction medium. This reduced demand foroxygen near the top of the reaction medium can be accomplished bycreating a vertical gradient in the concentration of the oxidizablecompound (e.g., para-xylene), where the minimum concentration ofoxidizable compound is located near the top of the reaction medium.

With respect to quantifying the oxidizable compound (e.g., para-xylene)concentration gradient, when the reaction medium is theoreticallypartitioned into 30 discrete horizontal slices of equal volume, anOC-max horizontal slice is identified as having the maximum oxidizablecompound concentration of all the 30 horizontal slices and an OC-minhorizontal slice is identified as having the minimum oxidizable compoundconcentration of the horizontal slices located above the OC-maxhorizontal slice. The oxidizable compound concentrations of thehorizontal slices are measured in the liquid phase on a time-averagedand volume-averaged mass fraction basis. It is preferred for the ratioof the oxidizable compound concentration of the OC-max horizontal sliceto the oxidizable compound concentration of the OC-min horizontal sliceto be greater than about 5:1, more preferably greater than about 10:1,still more preferably greater than about 20:1, and most preferably inthe range of from 40:1 to 1000:1.

Typically, the OC-max horizontal slice will be located near the bottomof the reaction medium, while the OC-min horizontal slice will belocated near the top of the reaction medium. Preferably, the OC-minhorizontal slice is one of the 5 upper-most horizontal slices of the 30discrete horizontal slices. Most preferably, the OC-min horizontal sliceis the upper-most one of the 30 discrete horizontal slices, asillustrated in FIG. 26. Preferably, the OC-max horizontal slice is oneof the 10 lower-most horizontal slices of the 30 discrete horizontalslices. Most preferably, the OC-max horizontal slice is one of the 5lower-most horizontal slices of the 30 discrete horizontal slices. Forexample, FIG. 26 illustrates the OC-max horizontal slice as the fifthhorizontal slice from the bottom of the reactor. It is preferred for thevertical spacing between the OC-min and OC-max horizontal slices to beat least about 2 W, where “W” is the maximum width of the reactionmedium. More preferably, the vertical spacing between the OC-min andOC-max horizontal slices is at least about 4 W, and most preferably atleast 6 W. Given a height “H” of the reaction medium, it is preferredfor the vertical spacing between the OC-min and OC-max horizontal slicesto be at least about 0.2 H, more preferably at least about 0.4 H, andmost preferably at least 0.6 H.

The time-averaged and volume-averaged oxidizable compound (e.g.,para-xylene) concentration in the liquid phase of the OC-min horizontalslice is preferably less than about 5,000 ppmw, more preferably lessthan about 2,000 ppmw, still more preferably less than about 400 ppmw,and most preferably in the range of from 1 ppmw to 100 ppmw. Thetime-averaged and volume-averaged oxidizable compound concentration inthe liquid phase of the OC-max horizontal slice is preferably in therange of from about 100 ppmw to about 10,000 ppmw, more preferably inthe range of from about 200 ppmw to about 5,000 ppmw, and mostpreferably in the range of from 500 ppmw to 3,000 ppmw.

Although it is preferred for the bubble column reactor to providevertical gradients in the concentration of the oxidizable compound, itis also preferred that the volume percent of the reaction medium havingan oxidizable compound concentration in the liquid phase above 1,000ppmw be minimized Preferably, the time-averaged volume percent of thereaction medium having an oxidizable compound concentration in theliquid phase above 1,000 ppmw is less than about 9 percent, morepreferably less than about 6 percent, and most preferably less than 3percent. Preferably, the time-averaged volume percent of the reactionmedium having an oxidizable compound concentration in the liquid phaseabove 2,500 ppmw is less than about 1.5 percent, more preferably lessthan about 1 percent, and most preferably less than 0.5 percent.Preferably, the time-averaged volume percent of the reaction mediumhaving an oxidizable compound concentration in the liquid phase above10,000 ppmw is less than about 0.3 percent, more preferably less thanabout 0.1 percent, and most preferably less than 0.03 percent.Preferably, the time-averaged volume percent of the reaction mediumhaving an oxidizable compound concentration in the liquid phase above25,000 ppmw is less than about 0.03 percent, more preferably less thanabout 0.015 percent, and most preferably less than 0.007 percent. Theinventors note that the volume of the reaction medium having theelevated levels of oxidizable compound need not lie in a singlecontiguous volume. At many times, the chaotic flow patterns in a bubblecolumn reaction vessel produce simultaneously two or more continuous butsegregated portions of the reaction medium having the elevated levels ofoxidizable compound. At each time used in the time averaging, all suchcontinuous but segregated volumes larger than 0.0001 volume percent ofthe total reaction medium are added together to determine the totalvolume having the elevated levels of oxidizable compound concentrationin the liquid phase.

In addition to the concentration gradients of oxygen and oxidizablecompound, discussed above, it is preferred for a temperature gradient toexist in the reaction medium. Referring again to FIG. 26, thistemperature gradient can be quantified in a manner similar to theconcentration gradients by theoretically partitioning the reactionmedium into 30 discrete horizontal slices of equal volume and measuringthe time-averaged and volume-averaged temperature of each slice. Thehorizontal slice with the lowest temperature out of the lowest 15horizontal slices can then be identified as the T-min horizontal slice,and the horizontal slice located above the T-min horizontal slice andhaving the maximum temperature of all the slices above the T-minhorizontal slice can then be identified as the “T-max horizontal slice.”It is preferred for the temperature of the T-max horizontal slice to beat least about 1° C. higher than the temperature of the T-min horizontalslice. More preferably the temperature of the T-max horizontal slice isin the range of from about 1.25 to about 12° C. higher than thetemperature of the T-min horizontal slice. Most preferably thetemperature of the T-max horizontal slice is in the range of from 2 to8° C. higher than the temperature of the T-min horizontal slice. Thetemperature of the T-max horizontal slice is preferably in the range offrom about 125 to about 200° C., more preferably in the range of fromabout 140 to about 180° C., and most preferably in the range of from 150to 170° C.

Typically, the T-max horizontal slice will be located near the center ofthe reaction medium, while the T-min horizontal slice will be locatednear the bottom of the reaction medium. Preferably, the T-min horizontalslice is one of the 10 lower-most horizontal slices of the 15 lowesthorizontal slices. Most preferably, the T-min horizontal slice is one ofthe 5 lower-most horizontal slices of the 15 lowest horizontal slices.For example, FIG. 26 illustrates the T-min horizontal slice as thesecond horizontal slice from the bottom of the reactor. Preferably, theT-max horizontal slice is one of the 20 middle horizontal slices of the30 discrete horizontal slices. Most preferably, the T-min horizontalslice is one of the 14 middle horizontal slices of the 30 discretehorizontal slices. For example, FIG. 26 illustrates the T-max horizontalslice as the twentieth horizontal slice from the bottom of the reactor(i.e., one of the middle 10 horizontal slices). It is preferred for thevertical spacing between the T-min and T-max horizontal slices to be atleast about 2 W, more preferably at least about 4 W, and most preferablyat least 6 W. It is preferred for the vertical spacing between the T-minand T-max horizontal slices to be at least about 0.2 H, more preferablyat least about 0.4 H, and most preferably at least 0.6 H.

As discussed above, when a vertical temperature gradient exists in thereaction medium, it can be advantageous to withdraw the reaction mediumat an elevated location where the temperature of reaction medium ishighest, especially when the withdrawn product is subjected to furtherdownstream processing at higher temperatures. Thus, when reaction medium36 is withdrawn from the reaction zone via one or more elevated outlets,as illustrated in FIGS. 15 and 16, it is preferred for the elevatedoutlet(s) to be located near the T-max horizontal slice. Preferably, theelevated outlet is located within 10 horizontal slices of the T-maxhorizontal slice, more preferably within 5 horizontal slices of theT-max horizontal slice, and most preferably within 2 horizontal slicesof the T-max horizontal slice.

It is now noted that many of the inventive features described herein canbe employed in multiple oxidation reactor systems—not just systemsemploying a single oxidation reactor. In addition, certain inventivefeatures described herein can be employed in mechanically-agitatedand/or flow-agitated oxidation reactors—not just bubble-agitatedreactors (i.e., bubble column reactors). For example, the inventors havediscovered certain advantages associated with staging/varying oxygenconcentration and/or oxygen consumption rate throughout the reactionmedium. The advantages realized by the staging of oxygenconcentration/consumption in the reaction medium can be realized whetherthe total volume of the reaction medium is contained in a single vesselor in multiple vessels. Further, the advantages realized by the stagingof oxygen concentration/consumption in the reaction medium can berealized whether the reaction vessel(s) is mechanically-agitated,flow-agitated, and/or bubble-agitated.

One way of quantifying the degree of staging of oxygen concentrationand/or consumption rate in a reaction medium is to compare two or moredistinct 20-percent continuous volumes of the reaction medium. These20-percent continuous volumes need not be defined by any particularshape. However, each 20-percent continuous volume must be formed of acontiguous volume of the reaction medium (i.e., each volume is“continuous”), and the 20-percent continuous volumes must not overlapone another (i.e., the volumes are “distinct”). These distinct20-percent continuous volumes can be located in the same reactor (FIG.29) or in multiple reactors. Referring now to FIG. 27, the bubble columnreactor is illustrated as containing a reaction medium that includes afirst distinct 20-percent continuous volume 37 and a second distinct20-percent continuous volume 39.

The staging of oxygen availability in the reaction medium can bequantified by referring to the 20-percent continuous volume of reactionmedium having the most abundant mole fraction of oxygen in the gas phaseand by referring to the 20-percent continuous volume of reaction mediumhaving the most depleted mole fraction of oxygen in the gas phase. Inthe gas phase of the distinct 20-percent continuous volume of thereaction medium containing the highest concentration of oxygen in thegas phase, the time-averaged and volume-averaged oxygen concentration,on a wet basis, is preferably in the range of from about 3 to about 18mole percent, more preferably in the range of from about 3.5 to about 14mole percent, and most preferably in the range of from 4 to 10 molepercent. In the gas phase of the distinct 20-percent continuous volumeof the reaction medium containing the lowest concentration of oxygen inthe gas phase, the time-averaged and volume-averaged oxygenconcentration, on a wet basis, is preferably in the range of from about0.3 to about 5 mole percent, more preferably in the range of from about0.6 to about 4 mole percent, and most preferably in the range of from0.9 to 3 mole percent. Furthermore, the ratio of the time-averaged andvolume-averaged oxygen concentration, on a wet basis, in the mostabundant 20-percent continuous volume of reaction medium compared to themost depleted 20-percent continuous volume of reaction medium ispreferably in the range of from about 1.5:1 to about 20:1, morepreferably in the range of from about 2:1 to about 12:1, and mostpreferably in the range of from 3:1 to 9:1.

The staging of oxygen consumption rate in the reaction medium can bequantified in terms of an oxygen-STR, initially described above.Oxygen-STR was previously describe in a global sense (i.e., from theperspective of the average oxygen-STR of the entire reaction medium);however, oxygen-STR may also be considered in a local sense (i.e., aportion of the reaction medium) in order to quantify staging of theoxygen consumption rate throughout the reaction medium.

The inventors have discovered that it is very useful to cause theoxygen-STR to vary throughout the reaction medium in general harmonywith the desirable gradients disclosed herein relating to pressure inthe reaction medium and to the mole fraction of molecular oxygen in thegas phase of the reaction medium. Thus, it is preferable that the ratioof the oxygen-STR of a first distinct 20-percent continuous volume ofthe reaction medium compared to the oxygen-STR of a second distinct20-percent continuous volume of the reaction medium be in the range offrom about 1.5:1 to about 20:1, more preferably in the range of fromabout 2:1 to about 12:1, and most preferably in the range of from 3:1 to9:1. In one embodiment the “first distinct 20-percent continuous volume”is located closer than the “second distinct 20-percent continuousvolume” to the location where molecular oxygen is initially introducedinto the reaction medium. These large gradients in oxygen-STR aredesirable whether the partial oxidation reaction medium is contained ina bubble column oxidation reactor or in any other type of reactionvessel in which gradients are created in pressure and/or mole fractionof molecular oxygen in the gas phase of the reaction medium (e.g., in amechanically agitated vessel having multiple, vertically disposedstirring zones achieved by using multiple impellers having strong radialflow, possibly augmented by generally horizontal baffle assemblies, withoxidant flow rising generally upwards from a feed near the lower portionof the reaction vessel, notwithstanding that considerable back-mixing ofoxidant flow may occur within each vertically disposed stirring zone andthat some back-mixing of oxidant flow may occur between adjacentvertically disposed stirring zones). That is, when a gradient exists inthe pressure and/or mole fraction of molecular oxygen in the gas phaseof the reaction medium, the inventors have discovered that it isdesirable to create a similar gradient in the chemical demand fordissolved oxygen by the means disclosed herein.

A preferred means of causing the local oxygen-STR to vary is bycontrolling the locations of feeding the oxidizable compound and bycontrolling the mixing of the liquid phase of the reaction medium tocontrol gradients in concentration of oxidizable compound according toother disclosures of the present invention. Other useful means ofcausing the local oxygen-STR to vary include causing variation inreaction activity by causing local temperature variation and by changingthe local mixture of catalyst and solvent components (e.g., byintroducing an additional gas to cause evaporative cooling in aparticular portion of the reaction medium and by adding a solvent streamcontaining a higher amount of water to decrease activity in a particularportion of the reaction medium).

When the oxidation reactor has a reactor-in-reactor configuration, asdescribed above with respect to FIGS. 12-14, it is preferred for theconcentration gradients, temperature gradients, and oxygen-STR gradientsdescribed herein with reference to FIGS. 26 and 27 to apply to theportion of the reaction medium located inside the external reactor andoutside the internal reactor (e.g., reaction medium 220 a in FIG. 12).

Referring again to FIGS. 1-27, oxidation is preferably carried out inthe bubble column reactor under conditions that are markedly different,according to preferred embodiments disclosed herein, than conventionaloxidation reactors. When the bubble column reactor is used to carry outthe liquid-phase partial oxidation of para-xylene to crude terephthalicacid (CTA) according to preferred embodiments disclosed herein, thespatial profiles of local reaction intensity, of local evaporationintensity, and of local temperature combined with the liquid flowpatterns within the reaction medium and the preferred, relatively lowoxidation temperatures contribute to the formation of CTA particleshaving unique and advantageous properties.

FIGS. 28A and 28B illustrate base CTA particles produced in accordancewith one embodiment of the present invention. FIG. 28A shows the baseCTA particles at 500 times magnification, while FIG. 28B zooms in on oneof the base CTA particles and shows that particle at 2,000 timesmagnification. As perhaps best illustrated in FIG. 28B, each base CTAparticle is typically formed of a large number of small, agglomeratedCTA subparticles, thereby giving the base CTA particle a relatively highsurface area, high porosity, low density, and good dissolvability.Unless otherwise specified, the various properties of the inventive CTA,described below, are measured using a representative sample of the CTA,where the representative sample weighs at least 1 gram and/or is formedof at least 10,000 individual CTA particles. The base CTA particlestypically have a mean particle size in the range of from about 20 toabout 150 microns, more preferably in the range of from about 30 toabout 120 microns, and most preferably in the range of from 40 to 90microns. The CTA subparticles typically have a mean particle size in therange of from about 0.5 to about 30 microns, more preferably from about1 to about 15 microns, and most preferably in the range of from 2 to 5microns. The relatively high surface area of the base CTA particlesillustrated in FIGS. 28A and 28B, can be quantified using aBraunauer-Emmett-Teller (BET) surface area measurement method.Preferably, the base CTA particles have an average BET surface of atleast about 0.6 meters squared per gram (m²/g). More preferably, thebase CTA particles have an average BET surface area in the range of fromabout 0.8 to about 4 m²/g. Most preferably, the base CTA particles havean average BET surface area in the range of from 0.9 to 2 m²/g. Thephysical properties (e.g., particle size, BET surface area, porosity,and dissolvability) of the base CTA particles formed by optimizedoxidation process of a preferred embodiment of the present inventionpermit purification of the CTA particles by more effective and/oreconomical methods, as described in further detail below with respect toFIG. 31.

The mean particle size values provided above were determined usingpolarized light microscopy and image analysis. The equipment employed inthe particle size analysis included a Nikon E800 optical microscope witha 4x Plan Flour N.A. 0.13 objective, a Spot RT™ digital camera, and apersonal computer running Image Pro Plus™ V4.5.0.19 image analysissoftware. The particle size analysis method included the following mainsteps: (1) dispersing the CTA powders in mineral oil; (2) preparing amicroscope slide/cover slip of the dispersion; (3) examining the slideusing polarized light microscopy (crossed polars condition—particlesappear as bright objects on black background); (4) capturing differentimages for each sample preparation (field size=3×2.25 mm; pixelsize=1.84 microns/pixel); (5) performing image analysis with Image ProPlus™ software; (6) exporting the particle measures to a spreadsheet;and (7) performing statistical characterization in the spreadsheet. Step(5) of “performing image analysis with Image Pro Plus™ software”included the substeps of: (a) setting the image threshold to detectwhite particles on dark background; (b) creating a binary image; (c)running a single-pass open filter to filter out pixel noise; (d)measuring all particles in the image; and (e) reporting the meandiameter measured for each particle. The Image Pro Plus™ softwaredefines mean diameter of individual particles as the number averagelength of diameters of a particle measured at 2 degree intervals andpassing through the particle's centroid. Step 7 of “performingstatistical characterization in the spreadsheet” comprises calculatingthe volume-weighted mean particle size as follows. The volume of each ofthe n particles in a sample is calculated as if it were spherical usingpi/6*d_(i)̂3; multiplying the volume of each particle times its diameterto find pi/6*d_(i)̂4; summing for all particles in the sample of thevalues of pi/6*d_(i)̂4; summing the volumes of all particles in thesample; and calculating the volume-weighted particle diameter as sum forall n particles in the sample of (pi/6*d_(i)̂4) divided by sum for all nparticles in the sample of (pi/6*d_(i)̂3). As used herein, “meanparticle size” refers to the volume-weighted mean particle sizedetermined according to the above-described test method; and it is alsoreferred to as D(4,3).

${D\left( {4,3} \right)} = \frac{\sum\limits_{i = 1}^{n}{\frac{\pi}{6}d_{i}^{4}}}{\sum\limits_{i = 1}^{n}{\frac{\pi}{6}d_{i}^{3}}}$

In addition, step 7 comprises finding the particle sizes for whichvarious fractions of the total sample volume are smaller. For example,D(v,0.1) is the particle size for which 10 percent of the total samplevolume is smaller and 90 percent is larger; D(v,0.5) is the particlesize for which one-half of the sample volume is larger and one-half issmaller; D(v,0.9) is the particle size for which 90 percent of the totalsample volume is smaller; and so on. In addition, step 7 comprisescalculating the value of D(v,0.9) minus D(v,0.1), which is hereindefined as the “particle size spread”; and step 7 comprises calculatingthe value of the particle size spread divided by D(4,3), which is hereindefined as the “particle size relative spread.”

Furthermore, it is preferable that the D(v,0.1) of the CTA particles asmeasured above be in the range from about 5 to about 65 microns, morepreferably in the range from about 15 to about 55 microns and mostpreferably in the range from 25 to 45 microns. It is preferable that theD(v,0.5) of the CTA particles as measured above be in the range fromabout 10 to about 90 microns, more preferably in the range from about 20to about 80 microns, and most preferably in the range from 30 to 70microns. It is preferable that the D(v,0.9) of the CTA particles asmeasured above be in the range from about 30 to about 150 microns, morepreferably in the range from about 40 to about 130 microns, and mostpreferably in the range from 50 to 110 microns. It is preferable thatthe particle size relative spread be in the range from about 0.5 toabout 2.0, more preferably in the range from about 0.6 to about 1.5, andmost preferably in the range from 0.7 to 1.3.

The BET surface area values provided above were measured on a

Micromeritics ASAP2000 (available from Micromeritics InstrumentCorporation of Norcross, Ga.). In the first step of the measurementprocess, a 2 to 4 gram of sample of the particles was weighed and driedunder vacuum at 50° C. The sample was then placed on the analysis gasmanifold and cooled to 77° K. A nitrogen adsorption isotherm wasmeasured at a minimum of 5 equilibrium pressures by exposing the sampleto known volumes of nitrogen gas and measuring the pressure decline. Theequilibrium pressures were appropriately in the range of P/P₀=0.01-0.20,where P is equilibrium pressure and P₀ is vapor pressure of liquidnitrogen at 77° K. The resulting isotherm was then plotted according tothe following BET equation:

$\frac{P}{V_{a}\left( {P_{o} - P} \right)} = {\frac{1}{V_{m}C} + {\frac{C - 1}{V_{m}C}\left( \frac{P}{P_{o}} \right)}}$

where V_(a) is volume of gas adsorbed by sample at P, V_(m) is volume ofgas required to cover the entire surface of the sample with a monolayerof gas, and C is a constant. From this plot, V_(m) and C weredetermined. V_(m) was then converted to a surface area using the crosssectional area of nitrogen at 77° K. by:

$A = {\sigma \frac{V_{m}}{RT}}$

where σ is cross sectional area of nitrogen at 77° K., T is 77° K., andR is the gas constant.

As alluded to above, CTA formed in accordance with one embodiment of thepresent invention exhibits superior dissolution properties versesconventional CTA made by other processes. This enhanced dissolution rateallows the inventive CTA to be purified by more efficient and/or moreeffective purification processes. The following description addressesthe manner in which the rate of dissolution of CTA can be quantified.

The rate of dissolution of a known amount of solids into a known amountof solvent in an agitated mixture can be measured by various protocols.As used herein, a measurement method called the “timed dissolution test”is defined as follows. An ambient pressure of about 0.1 megapascal isused throughout the timed dissolution test. The ambient temperature usedthroughout the timed dissolution test is about 22° C. Furthermore, thesolids, solvent and all dissolution apparatus are fully equilibratedthermally at this temperature before beginning testing, and there is noappreciable heating or cooling of the beaker or its contents during thedissolution time period. A solvent portion of fresh,

HPLC analytical grade of tetrahydrofuran (>99.9 percent purity),hereafter THF, measuring 250 grams is placed into a cleaned KIMAX tallform 400 milliliter glass beaker (Kimble® part number 14020,Kimble/Kontes, Vineland, N.J.), which is uninsulated, smooth-sided, andgenerally cylindrical in form. A Teflon-coated magnetic stirring bar(VWR part number 58948-230, about 1-inch long with ⅜-inch diameter,octagonal cross section, VWR International, West Chester, Pa. 19380) isplaced in the beaker, where it naturally settles to the bottom. Thesample is stirred using a Variomag® multipoint 15 magnetic stirrer (H&PLabortechnik AG, Oberschleissheim, Germany) magnetic stirrer at asetting of 800 revolutions per minute. This stirring begins no more than5 minutes before the addition of solids and continues steadily for atleast 30 minutes after adding the solids. A solid sample of crude orpurified TPA particulates amounting to 250 milligrams is weighed into anon-sticking sample weighing pan. At a starting time designated as t=0,the weighed solids are poured all at once into the stirred THF, and atimer is started simultaneously. Properly done, the THF very rapidlywets the solids and forms a dilute, well-agitated slurry within 5seconds. Subsequently, samples of this mixture are obtained at thefollowing times, measured in minutes from t=0: 0.08, 0.25, 0.50, 0.75,1.00, 1.50, 2.00, 2.50, 3.00, 4.00, 5.00, 6.00, 8.00, 10.00, 15.00, and30.00. Each small sample is withdrawn from the dilute, well-agitatedmixture using a new, disposable syringe (Becton, Dickinson and Co, 5milliliter, REF 30163, Franklin Lakes, N.J. 07417) Immediately uponwithdrawal from the beaker, approximately 2 milliliters of clear liquidsample is rapidly discharged through a new, unused syringe filter (25 mmdiameter, 0.45 micron, Gelman GHP Acrodisc GF®, Pall Corporation, EastHills, N.Y. 11548) into a new, labeled glass sample vial. The durationof each syringe filling, filter placement, and discharging into a samplevial is correctly less than about 5 seconds, and this interval isappropriately started and ended within about 3 seconds either side ofeach target sampling time. Within about five minutes of each filling,the sample vials are capped shut and maintained at approximatelyconstant temperature until performing the following chemical analysis.After the final sample is taken at a time of 30 minutes past t=0, allsixteen samples are analyzed for the amount of dissolved TPA using aHPLC-DAD method generally as described elsewhere within this disclosure.However, in the present test, the calibration standards and the resultsreported are both based upon milligrams of dissolved TPA per gram of THFsolvent (hereafter “ppm in THF”). For example, if all of the 250milligrams of solids were very pure TPA and if this entire amount fullydissolved in the 250 grams of THF solvent before a particular samplewere taken, the correctly measured concentration would be about 1,000ppm in THF.

When CTA according to the present invention is subjected to the timeddissolution test described above, it is preferred that a sample taken atone minute past t=0 dissolves to a concentration of at least about 500ppm in THF, more preferably to at least 600 ppm in THF. For a sampletaken at two minutes past t=0, it is preferred that CTA according to thecurrent invention will dissolve to a concentration of at least about 700ppm in THF, more preferably to at least 750 ppm in THF. For a sampletaken at four minutes past t=0, it is preferred that CTA according tothe current invention will dissolve to a concentration of at least about840 ppm in THF, more preferably to at least 880 ppm in THF.

The inventors have found that a relatively simple negative exponentialgrowth model is useful to describe the time dependence of the entiredata set from a complete timed dissolution test, notwithstanding thecomplexity of the particulate samples and of the dissolution process.The form of the equation, hereinafter the “timed dissolution model,” isas follows:

S=A+B*(1−exp(−C*t)), where

-   -   t=time in units of minutes;    -   S=solubility, in units of ppm in THF, at time t;    -   exp=exponential function in the base of the natural logarithm of        2;    -   A, B=regressed constants in units of ppm in THF, where A relates        mostly to the rapid dissolution of the smaller particles at very        short times, and where the sum of A+B relates mostly to the        total amount of dissolution near the end of the specified        testing period; and    -   C=a regressed time constant in units of reciprocal minutes.

The regressed constants are adjusted to minimize the sum of the squaresof the errors between the actual data points and the corresponding modelvalues, which method is commonly called a “least squares” fit. Apreferred software package for executing this data regression is JMPRelease 5.1.2 (SAS Institute Inc., JMP Software, SAS Campus Drive, Cary,N.C. 27513).

When CTA according to the present invention is tested with the timeddissolution test and fitted to the timed dissolution model describedabove, it is preferred for the CTA to have a time constant “C” greaterthan about 0.5 reciprocal minutes, more preferably greater than about0.6 reciprocal minutes, and most preferably greater than 0.7 reciprocalminutes.

FIGS. 29A and 29B illustrate a conventional CTA particle made by aconventional high-temperature oxidation process in a continuous stirredtank reactor (CSTR). FIG. 29A shows the conventional CTA particle at 500times magnification, while FIG. 29B zooms in and shows the CTA particleat 2,000 times magnification. A visual comparison of the inventive CTAparticles illustrated in FIGS. 28A and 28B and the conventional CTAparticle illustrated in FIGS. 29A and 29B shows that the conventionalCTA particle has a higher density, lower surface area, lower porosity,and larger particle size than the inventive CTA particles. In fact, theconventional CTA represented in FIGS. 29A and 29B has a mean particlesize of about 205 microns and a BET surface area of about 0.57 m²/g.

FIG. 30 illustrates a conventional process for making purifiedterephthalic acid (PTA). In the conventional PTA process, para-xylene ispartially oxidized in a mechanically agitated high temperature oxidationreactor 700. A slurry comprising CTA is withdrawn from reactor 700 andthen purified in a purification system 702. The PTA product ofpurification system 702 is introduced into a separation system 706 forseparation and drying of the PTA particles. Purification system 702represents a large portion of the costs associated with producing PTAparticles by conventional methods. Purification system 702 generallyincludes a water addition/exchange system 708, a dissolution system 710,a hydrogenation system 712, and three separate crystallization vessels704 a,b,c. In water addition/exchange system 708, a substantial portionof the mother liquor is displaced with water. After water addition, thewater/CTA slurry is introduced into the dissolution system 710 where thewater/CTA mixture is heated until the CTA particles fully dissolve inthe water. After CTA dissolution, the CTA-in-water solution is subjectedto hydrogenation in hydrogenation system 712. The hydrogenated effluentfrom hydrogenation system 712 is then subjected to three crystallizationsteps in crystallization vessels 704 a,b,c, followed by PTA separationin separation system 706.

FIG. 31 illustrates an improved process for producing PTA employing aoxidation reactor system comprising a primary oxidation reactor 800 aand a secondary oxidation reactor 800 b. In the configurationillustrated in FIG. 31, an initial slurry is produced from primaryoxidation reactor 800 a and is thereafter subjected to purification in apurification system 802, of which secondary oxidation reactor 800 b is apart. The initial slurry withdrawn from primary oxidation reactor 800 apreferably comprises solid CTA particles and a liquid mother liquor.Typically, the initial slurry contains in the range of from about 10 toabout 50 weight percent solid CTA particles, with the balance beingliquid mother liquor. The solid CTA particles present in the initialslurry withdrawn from primary oxidation reactor 800 a typically containat least about 400 ppmw of 4-carboxybenzaldehyde (4-CBA), more typicallyat least about 800 ppmw of 4-CBA, and most typically in the range offrom 1,000 to 15,000 ppmw of 4-CBA.

Purification system 802 receives the initial slurry withdrawn fromprimary oxidation reactor 800 a and reduces the concentration of 4-CBAand other impurities present in the CTA. A purer/purified slurry isproduced from purification system 802 and is subjected to separation anddrying in a separation system 804 to thereby produce purer solidterephthalic acid particles comprising less than about 400 ppmw of4-CBA, more preferably less than about 250 ppmw of 4-CBA, and mostpreferably in the range of from 10 to 200 ppmw of 4-CBA.

Purification system 802 includes secondary oxidation reactor 800 b, aliquor exchange system 806, a digester 808, and a single crystallizer810. In secondary oxidation reactor 800 b, the initial slurry issubjected to oxidation at a temperature and pressure that areapproximately equal to the temperature and pressure in primary oxidationreactor 800 a. In liquor exchange system 806, at least about 50 weightpercent of the mother liquor present in the slurry withdrawn fromsecondary oxidation reactor 800 b is replaced with a fresh replacementsolvent to thereby provide a solvent-exchanged slurry comprising CTAparticles and the replacement solvent. The solvent-exchanged slurryexiting liquor exchange system 806 is introduced into digester 808. Indigester 808, a further oxidation reaction is preformed at slightlyhigher temperatures than were used in primary oxidation reactor 800 a.

As discussed above, the high surface area, small particle size, and lowdensity of the CTA particles produced in primary oxidation reactor 800 acause certain impurities trapped in the CTA particles to becomeavailable for oxidation in digester 808 without requiring completedissolution of the CTA particles in digester 808. Thus, the temperaturein digester 808 can be lower than many similar prior art processes. Thefurther oxidation carried out in digester 808 preferably reduces theconcentration of 4-CBA in the CTA by at least 200 ppmw, more preferablyat least about 400 ppmw, and most preferably in the range of from 600 to6,000 ppmw. Preferably, the digestion temperature in digester 808 is atleast about 10° C. higher than the primary oxidation temperature inreactor 800 a, more preferably about 20 to about 80° C. higher than theprimary oxidation temperature in reactor 800 a, and most preferably 30to 50° C. higher than the primary oxidation temperature in reactor 800a. The digestion temperature is preferably in the range of from about160 to about 240° C., more preferably in the range of from about 180 toabout 220° C. and most preferably in the range of from 190 to 210° C.The purified product from digester 808 requires only a singlecrystallization step in crystallizer 810 prior to separation inseparation system 804. Suitable secondary oxidation/digestion techniquesare discussed in further detail in U.S. Pat. App. Pub. No. 2005/0065373,the entire disclosure of which is expressly incorporated herein byreference.

Terephthalic acid (e.g., PTA) produced by the system illustrated in FIG.31 is preferably formed of PTA particles having a mean particle size ofat least about 40 microns, more preferably in the range of from about 50to about 2,000 microns, and most preferably in the range of from 60 to200 microns. The PTA particles preferably have an average BET surfacearea less than about 0.25 m²/g, more preferably in the range of fromabout 0.005 to about 0.2 m²/g, and most preferably in the range of from0.01 to 0.18 m²/g. PTA produced by the system illustrated in FIG. 31 issuitable for use as a feedstock in the making of PET. Typically, PET ismade via esterification of terephthalic acid with ethylene glycol,followed by polycondensation. Preferably, terephthalic acid produced byan embodiment of the present invention is employed as a feed to the pipereactor PET process described in U.S. patent application Ser. No.10/013,318, filed Dec. 7, 2001, the entire disclosure of which isincorporated herein by reference.

CTA particles with the preferred morphology disclosed herein areparticularly useful in the above-described oxidative digestion processfor reduction of 4-CBA content. In addition, these preferred CTAparticles provide advantages in a wide range of other post-processesinvolving dissolution and/or chemical reaction of the particles. Theseadditional post-processes include, but are not limited too, reactionwith at least one hydroxyl-containing compound to form ester compounds,especially the reaction of CTA with methanol to form dimethylterephthalate and impurity esters; reaction with at least one diol toform ester monomer and/or polymer compounds, especially the reaction ofCTA with ethylene glycol to form polyethylene terephthalate (PET); andfull or partial dissolution in solvents, including, but not limited too,water, acetic acid, and N-methyl-2-pyrrolidone, which may includefurther processing, including, but not limited too, reprecipitation of amore pure terephthalic acid and/or selective chemical reduction ofcarbonyl groups other than carboxylic acid groups. Notably included isthe substantial dissolution of CTA in a solvent comprising water coupledwith partial hydrogenation that reduces the amount of aldehydes,especially 4-CBA, fluorenones, phenones, and/or anthraquinones.

In accordance with one embodiment of the present invention, there isprovided a process for partially oxidizing an oxidizable aromaticcompound to one or more types of aromatic carboxylic acid wherein thepurity of the solvent portion of the feed (i.e., the “solvent feed”) andthe purity of the oxidizable compound portion of the feed (i.e., the“oxidizable compound feed”) are controlled within certain rangesspecified below. Along with other embodiments of the present invention,this enables the purity of the liquid phase and, if present, the solidphase and the combined slurry (i.e., solid plus liquid) phase of thereaction medium to be controlled in certain preferred ranges, outlinedbelow.

With respect to the solvent feed, it is known to oxidize an oxidizablearomatic compound(s) to produce an aromatic carboxylic acid wherein thesolvent feed introduced into the reaction medium is a mixture ofanalytical-purity acetic acid and water, as is often employed atlaboratory scale and pilot scale. Likewise, it is known to conduct theoxidation of oxidizable aromatic compound to aromatic carboxylic acidwherein the solvent leaving the reaction medium is separated from theproduced aromatic carboxylic acid and then recycled back to the reactionmedium as feed solvent, primarily for reasons of manufacturing cost.This solvent recycling causes certain feed impurities and processby-products to accumulate over time in the recycled solvent. Variousmeans are known in the art to help purify recycled solvent beforere-introduction into the reaction medium. Generally, a higher degree ofpurification of the recycled solvent leads to significantly highermanufacturing cost than does a lower degree of purification by similarmeans. One embodiment of the present invention relates to understandingand defining the preferred ranges of a large number of impurities withinthe solvent feed, many of which were heretofore thought largely benign,in order to find an optimal balance between overall manufacturing costand overall product purity.

“Recycled solvent feed” is defined herein as solvent feed that waspreviously part of a reaction medium subjected to oxidation in anoxidation zone/reactor and exited the oxidation zone/reactor as part ofthe crude liquid and/or slurry product. For example, recycled solventfeed to a partial oxidation reaction medium for oxidizing para-xylene toform TPA is solvent that originally formed part of the partial oxidationreaction medium, was removed from the reaction medium as a liquid phaseof a TPA slurry, was separated away from most solid TPA mass, and wasthen returned to the partial oxidation reaction medium. As describedabove, such recycled solvent feed is prone to accumulate all manner ofundesirable impurities unless specific auxiliary process steps areprovided for solvent purification, at considerable capital and operatingcost. For economic reasons, it is preferable that at least about 20weight percent of the solvent feed to the reaction medium of the presentinvention is recycled solvent, more preferably at least about 40 weightpercent, still more preferably at least about 80 weight percent, andmost preferably at least 90 weight percent. For reasons of solventinventory and of on-stream time in a manufacturing unit, it ispreferable that portions of recycled solvent pass through reactionmedium at least once per day of operation, more preferably at least onceper day for at least seven consecutive days of operation, and mostpreferably at least once per day for at least 30 consecutive days ofoperation.

The inventors have discovered that, for reasons of reaction activity andfor consideration of metallic impurities left in the oxidation product,the concentrations of selected multivalent metals within the recycledsolvent feed are preferably in ranges specified immediately below. Theconcentration of iron in recycled solvent is preferably below about 150ppmw, more preferably below about 40 ppmw, and most preferably between 0and 8 ppmw. The concentration of nickel in recycled solvent ispreferably below about 150 ppmw, more preferably below about 40 ppmw,and most preferably between 0 and 8 ppmw. The concentration of chromiumin recycled solvent is preferably below about 150 ppmw, more preferablybelow about 40 ppmw, and most preferably between 0 and 8 ppmw. Theconcentration of molybdenum in recycled solvent is preferably belowabout 75 ppmw, more preferably below about 20 ppmw, and most preferablybetween 0 and 4 ppmw. The concentration of titanium in recycled solventis preferably below about 75 ppmw, more preferably below about 20 ppmw,and most preferably between 0 and 4 ppmw. The concentration of copper inrecycled solvent is preferably below about 20 ppmw, more preferablybelow about 4 ppmw, and most preferably between 0 and 1 ppmw. Othermetallic impurities are also typically present in recycled solvent,generally varying at lower levels in proportion to one or more of theabove listed metals. Controlling the above listed metals in thepreferred ranges will keep other metallic impurities at suitable levels.

These metals can arise as impurities in any of the incoming processfeeds (e.g., in incoming oxidizable compound, solvent, oxidant, andcatalyst compounds). Alternatively, the metals can arise as corrosionproducts from any of the process units contacting reaction medium and/orcontacting recycled solvent. The means for controlling the metals in thedisclosed concentration ranges include the appropriate specification andmonitoring of the purity of various feeds and the appropriate usage ofmaterials of construction, including, but not limited to, manycommercial grades of titanium and of stainless steels including thosegrades known as duplex stainless steels and high molybdenum stainlesssteels.

The inventors have also discovered preferred ranges for selectedaromatic compounds in the recycled solvent. These include bothprecipitated and dissolved aromatic compounds within the recycledsolvent.

Surprisingly, even precipitated product (e.g., TPA) from a partialoxidation of para-xylene, is a contaminant to be managed in recycledsolvent. Because there are surprisingly preferred ranges for the levelsof solids within the reaction medium, any precipitated product in thesolvent feed directly subtracts from the amount of oxidizable compoundthat can be fed in concert.

Furthermore, feeding precipitated TPA solids in the recycled solvent atelevated levels has been discovered to affect adversely the character ofthe particles formed within a precipitating oxidation medium, leading toundesirable character in downstream operations (e.g., productfiltration, solvent washing, oxidative digestion of crude product,dissolution of crude product for further processing, and so on). Anotherundesirable characteristic of precipitated solids in the recycle solventfeed is that these often contain very high levels of precipitatedimpurities, as compared to impurity concentrations in the bulk of thesolids within the TPA slurries from which much of the recycled solventis obtained. Possibly, the elevated levels of impurities observed insolids suspended in recycled solvent may relate to nucleation times forprecipitation of certain impurities from the recycled solvent and/or tocooling of the recycled solvent, whether intentional or due to ambientlosses. For example, concentrations of highly-colored and undesirable2,6-dicarboxyfluorenone have been observed at far higher levels insolids present in recycled solvent at 80° C. than are observed in TPAsolids separated from recycled solvent at 160° C. Similarly,concentrations of isophthalic acid have been observed at much higherlevels in solids present in recycled solvent compared to levels observedin TPA solids from the reaction medium. Exactly how specificprecipitated impurities entrained within recycled solvent behave whenre-introduced to the reaction medium appears to vary. This dependsperhaps upon the relative solubility of the impurity within the liquidphase of the reaction medium, perhaps upon how the precipitated impurityis layered within the precipitated solids, and perhaps upon the localrate of TPA precipitation where the solid first re-enters the reactionmedium. Thus, the inventors have found it useful to control the level ofcertain impurities in the recycled solvent, as disclosed below, withoutrespect to whether these impurities are present in the recycled solventin dissolved form or are entrained particulates therein.

The amount of precipitated solids present in recycled solvent isdetermined by a gravimetric method as follows. A representative sampleis withdrawn from the solvent supply to the reaction medium while thesolvent is flowing in a conduit toward the reaction medium. A usefulsample size is about 100 grams captured in a glass container havingabout 250 milliliters of internal volume. Before being released toatmospheric pressure, but while continuously flowing toward the samplecontainer, the recycled solvent is cooled to less than 100° C.; thiscooling is in order to limit solvent evaporation during the shortinterval before being sealed closed in the glass container. After thesample is captured at atmospheric pressure, the glass container issealed closed immediately. Then the sample is allowed to cool to about20° C. while surrounded by air at about 20° C. and without forcedconvection. After reaching about 20° C., the sample is held at thiscondition for at least about 2 hours. Then, the sealed container isshaken vigorously until a visibly uniform distribution of solids isobtained Immediately thereafter, a magnetic stirrer bar is added to thesample container and rotated at sufficient speed to maintain effectivelyuniform distribution of solids. A 10 milliliter aliquot of the mixedliquid with suspended solids is withdrawn by pipette and weighed. Thenthe bulk of the liquid phase from this aliquot is separated by vacuumfiltration, still at about 20° C. and effectively without loss ofsolids. The moist solids filtered from this aliquot are then dried,effectively without sublimation of solids, and these dried solids areweighed. The ratio of the weight of the dried solids to the weight ofthe original aliquot of slurry is the fraction of solids, typicallyexpressed as a percentage and referred to herein as the amount of“precipitated solids at 20° C.” in the solvent feed.

The inventors have discovered that aromatic compounds dissolved in theliquid phase of the reaction medium and comprising aromatic carboxylicacids lacking non-aromatic hydrocarbyl groups (e.g., isophthalic acid,benzoic acid, phthalic acid, 2,5,4′-tricarboxybiphenyl) are surprisinglypernicious components. Although these compounds are much reduced inchemical activity in the subject reaction medium compared to oxidizablecompounds having non-aromatic hydrocarbyl groups, the inventors havediscovered that these compounds nonetheless undergo numerous detrimentalreactions. Thus, it is advantageous to control the content of thesecompounds in preferred ranges in the liquid phase of the reactionmedium. This leads to preferred ranges of select compounds in recycledsolvent feed and also to preferred ranges of select precursors in theoxidizable aromatic compound feed.

For example, in the liquid-phase partial oxidation of para-xylene toterephthalic acid (TPA), the inventors have discovered that thehighly-colored and undesirable impurity 2,7-dicarboxyfluorenone(2,7-DCF) is virtually undetectable in the reaction medium and productoff-take when meta-substituted aromatic compounds are at very low levelsin the reaction medium. The inventors have discovered that whenisophthalic acid impurity is present at increasing levels in the solventfeed, the formation of 2,7-DCF rises in almost direct proportion. Theinventors have also discovered that when meta-xylene impurity is presentin the feed of para-xylene, the formation of 2,7-DCF again rises almostin direct proportion. Furthermore, even if the solvent feed andoxidizable compound feed are devoid of meta-substituted aromaticcompounds, the inventors have discovered that some isophthalic acid isformed during a typical partial oxidation of very pure para-xylene,particularly when benzoic acid is present in the liquid phase of thereaction medium. This self-generated isophthalic acid may, owing to itsgreater solubility than TPA in solvent comprising acetic acid and water,build up over time in commercial units employing recycled solvent. Thus,the amount of isophthalic acid within solvent feed, the amount ofmeta-xylene within oxidizable aromatic compound feed, and the rate ofself-creation of isophthalic acid within the reaction medium are allappropriately considered in balance with each other and in balance withany reactions that consume isophthalic acid. Isophthalic acid has beendiscovered to undergo additional consumptive reactions besides theformation of 2,7-DCF, as are disclosed below. In addition, the inventorshave discovered that there are other issues to consider when settingappropriate ranges for the meta-substituted aromatic species in thepartial oxidation of para-xylene to TPA. Other highly-colored andundesirable impurities, such as 2,6-dicarboxyfluorenone (2,6-DCF),appear to relate greatly to dissolved, para-substituted aromaticspecies, which are always present with para-xylene feed to aliquid-phase oxidation. Thus, the suppression of 2,7-DCF is bestconsidered in perspective with the level of other colored impuritiesbeing produced.

For example, in the liquid-phase partial oxidation of para-xylene toTPA, the inventors have discovered that the formation of trimelliticacid rises as the levels isophthalic acid and phthalic acid rise withinthe reaction medium. Trimellitic acid is a tri-functional carboxylicacid leading to branching of polymer chains during production of PETfrom TPA. In many PET applications, branching levels must be controlledto low levels and hence trimellitic acid must be controlled to lowlevels in purified TPA. Besides leading to trimellitic acid, thepresence of meta-substituted and ortho-substituted species in thereaction medium also give rise to other tricarboxylic acids (e.g.,1,3,5-tricarboxybenzene). Furthermore, the increased presence oftricarboxylic acids in the reaction medium increases the amount oftetracarboxylic acid formation (e.g., 1,2,4,5-tetracarboxybenzene).Controlling the summed production of all aromatic carboxylic acidshaving more than two carboxylic acid groups is one factor in setting thepreferred levels of meta-substituted and ortho-substituted species inthe recycled solvent feed, in the oxidizable compound feed, and in thereaction medium according to the present invention.

For example, in the liquid-phase partial oxidation of para-xylene toTPA, the inventors have discovered that increased levels in the liquidphase of the reaction medium of several dissolved aromatic carboxylicacids lacking non-aromatic hydrocarbyl groups leads directly to theincreased production of carbon monoxide and carbon dioxide. Thisincreased production of carbon oxides represents a yield loss on bothoxidant and on oxidizable compound, the later since many of theco-produced aromatic carboxylic acids, which on the one hand may beviewed as impurities, on the other hand also have commercial value.Thus, appropriate removal of relatively soluble carboxylic acids lackingnon-aromatic hydrocarbyl groups from recycle solvent has an economicvalue in preventing yield loss of oxidizable aromatic compound and ofoxidant, in addition to suppressing the generation of highly undesirableimpurities such as various fluorenones and trimellitic acid.

For example, in the liquid-phase partial oxidation of para-xylene toTPA, the inventors have discovered that formation of2,5,4′-tricarboxybiphenyl is seemingly unavoidable. The2,5,4′-tricarboxybiphenyl is an aromatic tricarboxylic acid formed bythe coupling of two aromatic rings, perhaps by the coupling of adissolved para-substituted aromatic species with an aryl radical,perhaps an aryl radical formed by decarboxylation or decarbonylation ofa para-substituted aromatic species. Fortunately, the2,5,4′-tricarboxybiphenyl is typically produced at lower levels thantrimellitic acid and does not usually lead to significantly increaseddifficulties with branching of polymer molecules during production ofPET. However, the inventors have discovered that elevated levels of2,5,4′-tricarboxybiphenyl in a reaction medium comprising oxidation ofalkyl aromatics according to preferred embodiments of the presentinvention lead to increased levels of highly-colored and undesirable2,6-DCF. The increased 2,6-DCF is possibly created from the2,5,4′-tricarboxybiphenyl by ring closure with loss of a water molecule,though the exact reaction mechanism is not known with certainty. If2,5,4′-tricarboxybiphenyl, which is more soluble in solvent comprisingacetic acid and water than is TPA, is allowed to build up too highwithin recycled solvent, conversion rates to 2,6-DCF can becomeunacceptably large.

For example, in the liquid-phase partial oxidation of para-xylene toTPA, the inventors have discovered that aromatic carboxylic acidslacking non-aromatic hydrocarbyl groups (e.g., isophthalic acid)generally lead to mild suppression of the chemical activity of thereaction medium when present in the liquid phase at sufficientconcentration.

For example, in the liquid-phase partial oxidation of para-xylene toTPA, the inventors have discovered that precipitation is very oftennon-ideal (i.e. non-equilibrium) with respect to the relativeconcentrations of different chemical species in the solid phase and inthe liquid phase. Perhaps, this is because the precipitation rate isvery fast at the space-time reaction rates preferred herein, leading tonon-ideal co-precipitation of impurities, or even occlusion. Thus, whenit is desired to limit the concentration of certain impurities (e.g.,trimellitic acid and 2,6-DCF) within crude TPA, owing to theconfiguration of downstream unit operations, it is preferable to controltheir concentration in solvent feed as well as their generation ratewithin the reaction medium.

For example, the inventors have discovered that benzophenone compounds(e.g., 4,4′-dicarboxybenzophenone and 2,5,4′-tricarboxybenzophenone)made during partial oxidation of para-xylene, have undesirable effectsin a PET reaction medium even though benzophenone compounds are not ashighly colored in TPA per se as are fluorenones and anthraquinones.Accordingly, it is desirable to limit the presence of benzophenones andselect precursors in recycled solvent and in oxidizable compound feed.Furthermore, the inventors have discovered that the presence of elevatedlevels of benzoic acid, whether admitted in recycled solvent or formedwithin the reaction medium, leads to elevated rates of production of4,4′-dicarboxybenzophenone.

In review, the inventors have discovered and sufficiently quantified asurprising array of reactions for aromatic compounds lackingnon-aromatic hydrocarbyl groups that are present in the liquid-phasepartial oxidation of para-xylene to TPA. Recapping just the single caseof benzoic acid, the inventors have discovered that increased levels ofbenzoic acid in the reaction medium of certain embodiments of thepresent invention lead to greatly increased production of the highlycolored and undesirable 9-fluorenone-2-carboxylic acid, to greatlyincreased levels of 4,4′-dicarboxybiphenyl, to increased levels of4,4′-dicarboxybenzophenone, to a mild suppression of chemical activityof the intended oxidation of para-xylene, and to increased levels ofcarbon oxides and attendant yield losses. The inventors have discoveredthat increased levels of benzoic acid in the reaction medium also leadto increased production of isophthalic acid and phthalic acid, thelevels of which are desirably controlled in low ranges according tosimilar aspects of the current invention. The number and importance ofreactions involving benzoic acid are perhaps even more surprising sincesome recent inventors contemplate using benzoic acid in place of aceticacid as a primary component of solvent (See, e.g., U.S. Pat. No.6,562,997). Additionally, the present inventors have observed thatbenzoic acid is self-generated during oxidation of para-xylene at ratesthat are quite important relative to its formation from impurities, suchas toluene and ethylbenzene, commonly found in oxidizable compound feedcomprising commercial-purity para-xylene.

On the other hand, the inventors have discovered little value fromadditional regulation of recycled solvent composition in regard to thepresence of oxidizable aromatic compound and in regard to aromaticreaction intermediates that both retain non-aromatic hydrocarbyl groupsand are also relatively soluble in the recycled solvent. In general,these compounds are either fed to or created within the reaction mediumat rates substantially greater than their presence in recycled solvent;and the consumption rate of these compounds within the reaction mediumis great enough, retaining one or more non-aromatic hydrocarbyl groups,to limit appropriately their build-up within recycled solvent. Forexample, during partial oxidation of para-xylene in a multi-phasereaction medium, para-xylene evaporates to a limited extent along withlarge quantities of solvent. When this evaporated solvent exits thereactor as part of the off-gas and is condensed for recovery as recycledsolvent, a substantial portion of the evaporated para-xylene condensestherein as well. It is not necessary to limit the concentration of thispara-xylene in recycled solvent. For example, if solvent is separatedfrom solids upon slurry exiting a para-xylene oxidation reaction medium,this recovered solvent will contain a similar concentration of dissolvedpara-toluic acid to that present at the point of removal from thereaction medium. Although it may be important to limit the standingconcentration of para-toluic acid within the liquid phase of thereaction medium, see below, it is not necessary to regulate separatelythe para-toluic acid in this portion of recycled solvent owing to itsrelatively good solubility and to its low mass flow rate relative to thecreation of para-toluic acid within the reaction medium. Similarly, theinventors have discovered little reason to limit the concentrations inrecycled solvent of aromatic compounds with methyl substituents (e.g.toluic acids), aromatic aldehydes (e.g., terephthaldehyde), of aromaticcompounds with hydroxy-methyl substituents (e.g., 4-hydroxymethylbenzoicacid), and of brominated aromatic compounds retaining at least onenon-aromatic hydrocarbyl group (e.g., alpha-bromo-para-toluic acid)below those inherently found in the liquid phase exiting from thereaction medium occurring in the partial oxidation of xylene accordingto preferred embodiments of the present invention. Surprisingly, theinventors have also discovered that it is also not necessary to regulatein recycled solvent the concentration of selected phenols intrinsicallyproduced during partial oxidation of xylene, for these compounds arecreated and destroyed within the reaction medium at rates much greaterthan their presence in recycled solvent. For example, the inventors havediscovered that 4-hydroxybenzoic acid has relatively small effects onchemical activity in the preferred embodiments of the present inventionwhen co-fed at rates of over 2 grams of 4-hydroxybenzoic acid per 1kilogram of para-xylene, far higher than the natural presence inrecycled solvent, despite being reported by others as a significantpoison in similar reaction medium (See, e.g., W. Partenheimer, CatalysisToday 23 (1995) p. 81).

Thus, there are numerous reactions and numerous considerations insetting the preferred ranges of various aromatic impurities in thesolvent feed as now disclosed. These discoveries are stated in terms ofthe aggregated weight average composition of all solvent streams beingfed to the reaction medium during the course of a set time period,preferably one day, more preferably one hour, and most preferably oneminute. For example, if one solvent feed flows substantiallycontinuously with a composition of 40 ppmw of isophthalic acid at a flowrate of 7 kilograms per minute, a second solvent feed flowssubstantially continuously with a composition of 2,000 ppmw ofisophthalic acid at a flow rate of 10 kilograms per minute, and thereare no other solvent feed streams entering the reaction medium, then theaggregated weight average composition of the solvent feed is calculatedas (40*7+2,000*10)/(7+10)=1,193 ppmw of isophthalic acid. It is notablethat the weight of any oxidizable compound feed or of any oxidant feedthat are perhaps commingled with the solvent feed before entering thereaction medium are not considered in calculating the aggregated weightaverage composition of the solvent feed.

Table 1, below, lists preferred values for certain components in thesolvent feed introduced into the reaction medium. The solvent feedcomponents listed in Table 1 are as follows: 4-carboxybenzaldehyde(4-CBA), 4,4′-dicarboxystilbene (4,4′-DCS), 2,6-dicarboxyanthraquinone(2,6-DCA), 2,6-dicarboxyfluorenone (2,6-DCF), 2,7-dicarboxyfluorenone(2,7-DCF), 3,5-dicarboxyfluorenone (3,5-DCF), 9-fluorenone-2-carboxylicacid (9F-2CA), 9-fluorenone-4-carboxylic acid (9F-4CA), totalfluorenones including other fluorenones not individually listed (totalfluorenones), 4,4′-dicarboxybiphenyl (4,4′-DCB),2,5,4′-tricarboxybiphenyl (2,5,4′-TCB), phthalic acid (PA), isophthalicacid (IPA), benzoic acid (BA), trimellitic acid (TMA),2,6-dicarboxybenzocoumarin (2,6-DCBC), 4,4′-dicarboxybenzil (4,4′-DCBZ),4,4′-dicarboxybenzophenone (4,4′-DCBP), 2,5,4′-tricarboxybenzophenone(2,5,4′-TCBP), terephthalic acid (TPA), precipitated solids at 20° C.,and total aromatic carboxylic acids lacking non-aromatic hydrocarbylgroups. Table 1, below provides the preferred amounts of theseimpurities in CTA produced according to an embodiment of the presentinvention.

TABLE 1 Components of Solvent Feed Introduced into Reaction MediumComponent Preferred More Preferred Most Preferred Identification Amt.(ppmw) Amt. (ppmw) Amt. (ppmw) 4-CBA <1,200 30-600 60-300 4,4′-DCS <3 <2 <1 2,6-DCA <6 0.1-3 0.2-1 2,6-DCF <20 0.1-10   0.5-5   2,7-DCF <100.1-5   0.5-2   3,5-DCF <10  <5 <2 9F-2CA <10 0.1-5   0.5-2   9F-4CA <5 <3 <1 Total fluorenones <40 <20 1-8  4,4′-DCB <45 <15 0.5-5  2,5,4′-TCB <45 0.1-15   0.5-5   PA <1,000 15-400 40-150 IPA 2,500  40-1,200 120-400 BA <4,500   50-1,500 150-500  TMA <1,000 15-40040-150 2,6-DCBC <40 <20 <5 4,4′-DCBZ <40 <20 <5 4,4′-DCBP <40 <20 <52,5,4′-TCBP <40 <20 0.5-5   TPA <9,000  200-6,000  400-2,000Precipitated <9,000  200-6,000  600-2,000 Solids at 20° C. TotalAromatic <18,000  300-9,000  450-3,000 Carboxylic Acids Lacking Non-Aromatic Hydrocarbyl Groups

Many other aromatic impurities are also typically present in recycledsolvent, generally varying at even lower levels and/or in proportion toone or more of the disclosed aromatic compounds. Methods for controllingthe disclosed aromatic compounds in the preferred ranges will typicallykeep other aromatic impurities at suitable levels.

When bromine is used within the reaction medium, a large number of ionicand organic forms of bromine are known to exist in a dynamicequilibrium. These various forms of bromine have different stabilitycharacteristics once leaving the reaction medium and passing throughvarious unit operations pertaining to recycled solvent. For example,alpha-bromo-para-toluic acid may persist as such at some conditions ormay rapidly hydrolyze at other conditions to form 4-hydroxymethylbenzoicacid and hydrogen bromide. In the present invention, it is preferablethat at least about 40 weight percent, more preferable that at leastabout 60 weight percent, and most preferable that at least about 80weight percent of the total mass of bromine present in the aggregatedsolvent feed to the reaction medium is in one or more of the followingchemical forms: ionic bromine, alpha-bromo-para-toluic acid, andbromoacetic acid.

Although the importance and value of controlling the aggregated weightaverage purity of solvent feed within the disclosed, desired ranges ofthe present invention has not heretofore been discovered and/ordisclosed, suitable means for controlling the solvent feed purity may beassembled from various methods already known in the art. First, anysolvent evaporated from the reaction medium is typically of suitablepurity providing that liquid or solids from the reaction medium are notentrained with the evaporated solvent. The feeding of reflux solventdroplets into the off-gas disengaging space above the reaction medium,as disclosed herein, appropriately limits such entrainment; and recycledsolvent of suitable purity with respect to aromatic compound can becondensed from such off-gas. Second, the more difficult and costlypurification of recycled solvent feed typically relates to solvent takenfrom the reaction medium in liquid form and to solvent that subsequentlycontacts the liquid and/or solid phases of the reaction medium withdrawnfrom the reaction vessel (e.g., recycled solvent obtained from a filterin which solids are concentrated and/or washed, recycled solventobtained from a centrifuge in which solids are concentrated and/orwashed, recycled solvent taken from a crystallization operation, and soon). However, means are also known in the art for effecting thenecessary purification of these recycled solvent streams using one ormore prior disclosures. With respect to controlling precipitated solidsin recycled solvent to be within the ranges specified, suitable controlmeans include, but are not limited to, gravimetric sedimentation,mechanical filtration using filter cloth on rotary belt filters androtary drum filters, mechanical filtration using stationary filtermedium within pressure vessels, hydro-cyclones, and centrifuges. Withrespect to controlling dissolved aromatic species in recycled solvent tobe within the ranges specified, the control means include, but are notlimited to, those disclosed in U.S. Pat. No. 4,939,297 and U.S. Pat.App. Pub. No. 2005-0038288, incorporated herein by reference. However,none of these prior inventions discovered and disclosed the preferredlevels of purity in the aggregated solvent feed as disclosed herein.Rather, these prior inventions merely provided means to purify selectedand partial streams of recycled solvent without deducing the presentinventive, optimal values of the composition of the aggregated weightaverage solvent feed to the reaction medium.

Turning now to the purity of the feed of oxidizable compound, it isknown that certain levels of isophthalic acid, phthalic acid, andbenzoic acid are present and tolerable at low levels in purified TPAused for polymer production. Moreover, it is known these species arerelatively more soluble in many solvents and may be advantageouslyremoved from purified TPA by crystallization processes. However, from anembodiment of the invention disclosed herein, it is now known thatcontrolling the level of several relatively soluble aromatic species,notably including isophthalic acid, phthalic acid, and benzoic acid, inthe liquid phase of the reaction medium is surprisingly important forcontrolling the level of polycyclic and colored aromatic compoundscreated in the reaction medium, for controlling compounds with more than2 carboxylic acid functions per molecule, for controlling reactionactivity within the partial oxidation reaction medium, and forcontrolling yield losses of oxidant and of aromatic compound.

It is known within the art that isophthalic acid, phthalic acid, andbenzoic acid are formed in the reaction medium as follows. Meta-Xylenefeed impurity oxidizes in good conversion and yield to IPA. Ortho-Xylenefeed impurity oxidizes in good conversion and yield to phthalic acid.Ethylbenzene and toluene feed impurities oxidize in good conversion andyield to benzoic acid. However, the inventors have observed thatsignificant amounts of isophthalic acid, phthalic acid, and benzoic acidare also formed within a reaction medium comprising para-xylene by meansother than oxidation of meta-xylene, ortho-xylene, ethylbenzene, andtoluene. These other intrinsic chemical routes possibly includedecarbonylation, decarboxylation, the re-organization of transitionstates, and addition of methyl and carbonyl radicals to aromatic rings.

In determining preferred ranges of impurities in the feed of oxidizablecompound, many factors are relevant. Any impurity in the feed is likelyto be a direct yield loss and a product purification cost if the purityrequirements of the oxidized product are sufficiently strict (e.g., in areaction medium for partial oxidation of para-xylene, toluene andethylbenzene typically found in commercial-purity para-xylene lead tobenzoic acid, and this benzoic acid is largely removed from mostcommercial TPA). When the partial oxidation product of a feed impurityparticipates in additional reactions, factors other than simple yieldloss and removal become appropriate when considering how much feedpurification cost to incur (e.g., in a reaction medium for partialoxidation of para-xylene, ethylbenzene leads to benzoic acid, andbenzoic acid subsequently leads to highly colored9-fluorenone-2-carboxylic acid, to isophthalic acid, to phthalic acid,and to increased carbon oxides, among others). When the reaction mediumself-generates additional amounts of an impurity by chemical mechanismsnot directly related to feed impurities, the analysis becomes still morecomplex (e.g., in a reaction medium for partial oxidation ofpara-xylene, benzoic acid is also self-generated from para-xyleneitself). In addition, the downstream processing of the crude oxidationproduct may affect the considerations for preferred feed purity. Forexample, the cost of removing to suitable levels a direct impurity(benzoic acid) and subsequent impurities (isophthalic acid, phthalicacid, 9-fluorenone-2-carboxylic acid, et al.) may be one and the same,may be different from each other, and may be different from therequirements of removing a largely unrelated impurity (e.g., incompleteoxidation product 4-CBA in the oxidation of para-xylene to TPA).

The following disclosed feed purity ranges for para-xylene are preferredwhere para-xylene is fed with solvent and oxidant to a reaction mediumfor partial oxidation to produce TPA. These ranges are more preferred inTPA production process having post-oxidation steps to remove fromreaction medium impurities other than oxidant and solvent (e.g.,catalyst metals). These ranges are still more preferred in TPAproduction processes that remove additional 4-CBA from CTA (e.g., byconversion of CTA to dimethyl terephthalate plus impurity esters andsubsequent separation of the methyl ester of 4-CBA by distillation, byoxidative digestion methods for converting 4-CBA to TPA, byhydrogenation methods for converting 4-CBA to para-toluic acid, which isthen separated by partial-crystallization methods). These ranges aremost preferred in TPA production processes that remove additional 4-CBAfrom CTA by oxidative digestion methods for converting 4-CBA to TPA.

Using new knowledge of preferred ranges of recycling aromatic compoundsand of the relative amounts of the aromatic compounds formed directlyfrom oxidation of feed impurities as compared to other intrinsicchemical routes, improved ranges for impurities have been discovered forimpure para-xylene being fed to a partial oxidation process for TPAproduction. Table 2 below provides preferred values for the amount ofmeta-xylene, ortho-xylene, and ethylbenzene +toluene in the para-xylenefeed expressed in parts per million by weight of para-xylene.

TABLE 2 Components of Impure para-xylene Feed Component Preferred MorePreferred Most Preferred Identification Amt. (ppmw) Amt. (ppmw) Amt.(ppmw) meta-xylene 20-800 50-600 100-400 ortho-xylene 10-300 20-200 30-100 ethylbenzene + 20-700 50-500 100-300 toluene* total 50-900100-800  200-700 *Specification for ethylbenzene + toluene is eachseparately and in sum

Those skilled in the art will now recognize the above impurities withinimpure para-xylene may have their greatest effect on the reaction mediumafter their partial oxidation products have accumulated in recycledsolvent. For example, feeding the upper amount of the most preferredrange of meta-xylene, 400 ppmw, will immediately produce about 200 ppmwof isophthalic acid within the liquid phase of the reaction medium whenoperating with about 33 weight percent solids in the reaction medium.This compares with an input from the upper amount of the most preferredrange for isophthalic acid in recycled solvent of 400 ppmw that, afterallowing for a typical solvent evaporation to cool the reaction medium,amounts to about 1,200 ppmw of isophthalic acid within the liquid phaseof the reaction medium. Thus, it is the accumulation of partialoxidation products over time within recycled solvent that represents thegreatest probable impact of the meta-xylene, ortho-xylene, ethylbenzene,and toluene impurities in the feed of impure para-xylene. Accordingly,the above ranges for impurities in impure para-xylene feed are preferredto be maintained for at least one-half of each day of operation of anypartial oxidation reaction medium in a particular manufacturing unit,more preferably for at least three-quarters of each day for at leastseven consecutive days of operation, and most preferably when themass-weighted averages of the impure para-xylene feed composition arewithin the preferred ranges for at least 30 consecutive days ofoperation.

Means for obtaining impure para-xylene of preferred purity are alreadyknown in the art and include, but are not limited to, distillation,partial crystallization methods at sub-ambient temperatures, andmolecular sieve methods using selective pore-size adsorption. However,the preferred ranges of purity specified herein are, at their high end,more demanding and expensive than characteristically practiced bycommercial suppliers of para-xylene; and yet at the low end, thepreferred ranges avoid overly costly purification of para-xylene forfeeding to a partial oxidation reaction medium by discovering anddisclosing where the combined effects of impurity self-generation frompara-xylene itself and of impurity consumptive reactions within thereaction medium become more important than the feed rates of impuritieswithin impure para-xylene.

When the xylene-containing feed stream contains selected impurities,such as ethyl-benzene and/or toluene, oxidation of these impurities cangenerate benzoic acid. As used herein, the term “impurity-generatedbenzoic acid” shall denote benzoic acid derived from any source otherthan xylene during xylene oxidation.

As disclosed herein, a portion of the benzoic acid produced duringxylene oxidation is derived from the xylene itself. This production ofbenzoic acid from xylene is distinctly in addition to any portion ofbenzoic acid production that may be impurity-generated benzoic acid.Without being bound by theory, it is believed that benzoic acid isderived from xylene within the reaction medium when various intermediateoxidation products of xylene spontaneously decarbonylate (carbonmonoxide loss) or decarboxylate (carbon dioxide loss) to thereby producearyl radicals. These aryl radicals can then abstract a hydrogen atomfrom one of many available sources in the reaction medium and produceself-generated benzoic acid. Whatever the chemical mechanism, the term“self-generated benzoic acid,” as used herein, shall denote benzoic acidderived from xylene during xylene oxidation.

As also disclosed herein, when para-xylene is oxidized to produceterephthalic acid (TPA), the production of self-generated benzoic acidcauses para-xylene yield loss and oxidant yield loss. In addition, thepresence of self-generated benzoic acid in the liquid phase of thereaction medium correlates with increases for many undesirable sidereactions, notably including generation of highly colored compoundscalled mono-carboxy-fluorenones. Self-generated benzoic acid alsocontributes to the undesirable accumulation of benzoic acid in recycledsolvent, which further elevates the concentration of benzoic acid in theliquid phase of the reaction medium. Thus, formation of self-generatedbenzoic acid is desirably minimized, but this is also appropriatelyconsidered simultaneously with impurity-generated benzoic acid, withfactors affecting consumption of benzoic acid, with factors pertainingto other issues of reaction selectivity, and with overall economics.

The inventors have discovered that the self-generation of benzoic acidcan be controlled to low levels by appropriate selection of, forexample, temperature, xylene distribution, and oxygen availabilitywithin the reaction medium during oxidation. Not wishing to be bound bytheory, lower temperatures and improved oxygen availability appear tosuppress the decarbonylation and/or decarboxylation rates, thus avoidingthe yield loss aspect of self-generated benzoic acid. Sufficient oxygenavailability appears to direct aryl radicals toward other more benignproducts, in particular hydroxybenzoic acids. Distribution of xylene inthe reaction medium may also affect the balance between aryl radicalconversion to benzoic acid or to hydroxybenzoic acids. Whatever thechemical mechanisms, the inventors have discovered reaction conditionsthat, although mild enough to reduce benzoic acid production, are severeenough to oxidize a high fraction of the hydroxybenzoic acid productionto carbon monoxide and/or carbon dioxide, which are easily removed fromthe oxidation product.

In a preferred embodiment of the present invention, the oxidationreactor is configured and operated in a manner such that the formationof self-generated benzoic acid is minimized and the oxidation ofhydroxybenzoic acids to carbon monoxide and/or carbon dioxide ismaximized When the oxidation reactor is employed to oxidize para-xyleneto terephthalic acid, it is preferred that para-xylene makes up at leastabout 50 weight percent of the total xylene in the feed streamintroduced into the reactor. More preferably, para-xylene makes up atleast about 75 weight percent of the total xylene in the feed stream.Still more preferably, para-xylene makes up at least 95 weight percentof the total xylene in the feed stream. Most preferably, para-xylenemakes up substantially all of the total xylene in the feed stream.

When the reactor is employed to oxidize para-xylene to terephthalicacid, it is preferred for the rate of production of terephthalic acid tobe maximized, while the rate of production of self-generated benzoicacid is minimized. Preferably, the ratio of the rate of production (byweight) of terephthalic acid to the rate of production (by weight) ofself-generated benzoic acid is at least about 500:1, more preferably atleast about 1,000:1, and most preferably at least 1,500:1. As will beseen below, the rate of production of self-generated benzoic acid ispreferably measured when the concentration of benzoic acid in the liquidphase of the reaction medium is below 2,000 ppmw, more preferably below1,000 ppmw, and most preferably below 500 ppmw, because these lowconcentrations suppress to suitably low rates reactions that convertbenzoic acid to other compounds.

Combining the self-generated benzoic acid and the impurity-generatedbenzoic acid, the ratio of the rate of production (by weight) ofterephthalic acid to the rate of production (by weight) of total(self-generated and impurity-generated) benzoic acid is preferably atleast about 400:1, more preferably at least about 700:1, and mostpreferably at least 1,100:1. As will be seen below, the summed rate ofproduction of self-generated benzoic acid plus impurity-generatedbenzoic acid is preferably measured when the concentration of benzoicacid in the liquid phase of the reaction medium is below 500 ppmw,because these low concentrations suppress to suitably low ratesreactions that convert benzoic acid to other compounds.

As disclosed herein, elevated concentrations of benzoic acid in theliquid phase of the reaction medium lead to increased formation of manyother aromatic compounds, several of which are noxious impurities inTPA; and, as disclosed herein, elevated concentrations of benzoic acidin the liquid phase of the reaction medium lead to increased formationof carbon oxide gases, the formation of which represents yield loss onoxidant and on aromatic compounds and/or solvent. Furthermore, it is nowdisclosed that the inventors have discovered a considerable portion ofthis increased formation of other aromatic compounds and of carbonoxides derives from reactions that convert some of the benzoic acidmolecules themselves, as contrasted to benzoic acid catalyzing otherreactions without itself being consumed. Accordingly, the “netgeneration of benzoic acid” is defined herein as the time-averagedweight of all benzoic acid exiting the reaction medium minus thetime-averaged weight of all benzoic acid entering the reaction mediumduring the same period of time. This net generation of benzoic acid isoften positive, driven by the formation rates of impurity-generatedbenzoic acid and of self-generated benzoic acid. However, the inventorshave discovered that the conversion rate of benzoic acid to carbonoxides, and to several other compounds, appears to increaseapproximately linearly as the concentration of benzoic acid is increasedin the liquid phase of the reaction medium, measured when other reactionconditions comprising temperature, oxygen availability, STR, andreaction activity are maintained appropriately constant. Thus, when theconcentration of benzoic acid in the liquid-phase of the reaction mediumis great enough, perhaps due to an elevated concentration of benzoicacid in recycled solvent, then the conversion of benzoic acid moleculesto other compounds, including carbon oxides, can become equal to orgreater than the chemical generation of new benzoic acid molecules. Inthis case, the net generation of benzoic acid can become balanced nearzero or even negative. The inventors have discovered that when the netgeneration of benzoic acid is positive, then the ratio of the rate ofproduction (by weight) of terephthalic acid in the reaction mediumcompared to the rate of net generation of benzoic acid in the reactionmedium is preferably above about 700:1, more preferably above about1,100:1, and most preferably above 4,000:1. The inventors havediscovered that when the net generation of benzoic acid is negative, theratio of the rate of production (by weight) of terephthalic acid in thereaction medium compared to the rate of net generation of benzoic acidin the reaction medium is preferably above about 200:(−1), morepreferably above about 1,000:(−1), and most preferably above 5,000:(−1).

The inventors have also discovered preferred ranges for the compositionof the slurry (liquid+solid) withdrawn from the reaction medium and forthe solid CTA portion of the slurry. The preferred slurry and thepreferred CTA compositions are surprisingly superior and useful. Forexample, purified TPA produced from this preferred CTA by oxidativedigestion has a sufficiently low level of total impurities and ofcolored impurities such that the purified TPA is suitable, withouthydrogenation of additional 4-CBA and/or colored impurities, for a widerange of applications in PET fibers and PET packaging applications. Forexample, the preferred slurry composition provides a liquid phase of thereaction medium that is relatively low in concentration of importantimpurities and this importantly reduces the creation of other even moreundesirable impurities as disclosed herein. In addition, the preferredslurry composition importantly aids the subsequent processing of liquidfrom the slurry to become suitably pure recycled solvent, according toother embodiments of the present invention.

CTA produced according to one embodiment of the present inventioncontains less impurities of selected types than CTA produce byconventional processes and apparatuses, notably those employing recycledsolvent. Impurities that may be present in CTA include the following:4-carboxybenzaldehyde (4-CBA), 4,4′-dicarboxystilbene (4,4′-DCS),2,6-dicarboxyanthraquinone (2,6-DCA), 2,6-dicarboxyfluorenone (2,6-DCF),2,7-dicarboxyfluorenone (2,7-DCF), 3,5-dicarboxyfluorenone (3,5-DCF),9-fluorenone-2-carboxylic acid (9F-2CA), 9-fluorenone-4-carboxylic acid(9F-4CA), total fluorenones including other fluorenones not individuallylisted (total fluorenones), 4,4′-dicarboxybiphenyl (4,4′-DCB),2,5,4′-tricarboxybiphenyl (2,5,4′-TCB), phthalic acid (PA), isophthalicacid (IPA), benzoic acid (BA), trimellitic acid (TMA), para-toluic acid(PTAC), 2,6-dicarboxybenzocoumarin (2,6-DCBC), 4,4′-dicarboxybenzil(4,4′-DCBZ), 4,4′-dicarboxybenzophenone (4,4′-DCBP),2,5,4′-tricarboxybenzophenone (2,5,4′-TCBP). Table 3, below provides thepreferred amounts of these impurities in CTA produced according to anembodiment of the present invention.

TABLE 3 CTA Impurities Impurity Preferred More Preferred Most PreferredIdentification Amt. (ppmw) Amt. (ppmw) Amt. (ppmw) 4-CBA <15,000 100-8,000  400-2,000 4,4′-DCS <12 <6 <3 2,6-DCA <9 <6 <2 2,6-DCF <1002-50 5-25 2,7-DCF <30 <15 <5 3,5-DCF <16 <8 <2 9F-2CA <16 <8 <4 9F-4CA<8 <4 <2 Total fluorenones <100 2-60 4-35 4,4′-DCB <64 1-32 2-82,5,4′-TCB <24 <12 <8 PA <200  3-100 5-50 IPA <800 10-400 20-200 BA <600 5-300 15-100 TMA <800 10-400 20-200 PTAC <2,000   10-1,000 50-5002,6-DCBC <64 <32 <8 4,4′-DCBZ <12 <8 <4 4,4′-DCBP <40 <30 <20 2,5,4′-TCBP <32 <16 <4

In addition, it is preferred for CTA produced according to an embodimentof the present invention to have reduced color content relative to CTAproduce by conventional processes and apparatuses, notably thoseemploying recycled solvent. Thus, it is preferred for CTA produced inaccordance to one embodiment of the present invention to have a percenttransmittance percent at 340 nanometers (nm) of at least about 25percent, more preferably of at least about 50 percent, and mostpreferably of at least 60 percent. It is further preferred for CTAproduced in accordance to one embodiment of the present invention tohave a percent transmittance percent at 400 nanometers (nm) of at leastabout 88 percent, more preferably of at least about 90 percent, and mostpreferably of at least 92 percent.

The test for percent transmittance provides a measure of the colored,light-absorbing impurities present within TPA or CTA. As used herein,the test refers to measurements done on a portion of a solution preparedby dissolving 2.00 grams of dry solid TPA or CTA in 20.0 milliliters ofdimethyl sulfoxide (DMSO), analytical grade or better. A portion of thissolution is then placed in a Hellma semi-micro flow cell, PN 176.700,which is made of quartz and has a light path of 1.0 cm and a volume of0.39 milliliters. (Hellma USA, 80 Skyline Drive, Plainview, N.Y. 11803).An Agilent 8453 Diode Array Spectrophotometer is used to measure thetransmittance of different wavelengths of light through this filled flowcell. (Agilent Technologies, 395 Page Mill Road, Palo Alto, Calif.94303). After appropriate correction for absorbance from the background,including but not limited to the cell and the solvent used, the percenttransmittance results, characterizing the fraction of incident lightthat is transmitted through the solution, are reported directly by themachine. Percent transmittance values at light wavelengths of 340nanometers and 400 nanometers are particularly useful for discriminatingpure TPA from many of the impurities typically found therein.

The preferred ranges of various aromatic impurities in the slurry(solid+liquid) phase of the reaction medium are provided below in Table4.

TABLE 4 Slurry Impurities Impurity Preferred More Preferred MostPreferred Identification Amt. (ppmw) Amt. (ppmw) Amt. (ppmw) 4-CBA<8,000 <5,000 <2,500 4,4′-DCS <4 <2 <1 2,6-DCA <6 <3 <1 2,6-DCF <70 2-404-20 2,7-DCF <12 <8 <4 3,5-DCF <12 <8 <4 9F-2CA <12 <8 <4 9F-4CA <8 <4<2 Total fluorenones <90 2-60 5-30 4,4′-DCB <64 1-16 2-4  2,5,4′-TCB <602-40 4-20 PA <3,000   25-1,500 75-500 IPA 9,000   75-4,500  225-1,500 BA<15,000  100-6,000  300-2,000 TMA <3,000   25-1,500 75-500 PTAC <8,000 100-4,000  200-2,000 4,4′-DCBZ <5 <4 <3 4,4′-DCBP <240 <160 <802,5,4′-TCBP <120 <80 <40

These preferred compositions for the slurry embody the preferredcomposition of the liquid phase of the reaction medium while usefullyavoiding experimental difficulties pertaining to precipitation ofadditional liquid phase components from the reaction medium into solidphase components during sampling from the reaction medium, separation ofliquids and solids, and shifting to analytical conditions.

Many other aromatic impurities are also typically present in the slurryphase of the reaction medium and in CTA of the reaction medium,generally varying at even lower levels and/or in proportion to one ormore of the disclosed aromatic compounds. Controlling the disclosedaromatic compounds in the preferred ranges will keep other aromaticimpurities at suitable levels. These advantaged compositions for theslurry phase in the reaction medium and for the solid CTA taken directlyfrom the slurry are enabled by operating with embodiments of theinvention disclosed herein for partial oxidation of para-xylene to TPA.

Measurement of the concentration of low level components in the solvent,recycled solvent, CTA, slurry from the reaction medium, and PTA areperformed using liquid chromatography methods. Two interchangeableembodiments are now described.

The method referred to herein as HPLC-DAD comprises high pressure liquidchromatography (HPLC) coupled with a diode array detector (DAD) toprovide separation and quantitation of various molecular species withina given sample. The instrument used in this measurement is a model 1100HPLC equipped with a DAD, provided by Agilent Technologies (Palo Alto,Calif.), though other suitable instruments are also commerciallyavailable and from other suppliers As is known in the art, both theelution time and the detector response are calibrated using knowncompounds present in known amounts, compounds and amounts that areappropriate to those occurring in actual unknown samples.

The method referred to herein as HPLC-MS comprises high pressure liquidchromatography (HPLC) coupled with mass spectrometry (MS) to provideseparation, identification, and quantitation of various molecularspecies within a given sample. The instruments used in this measurementis an Alliance HPLC and ZQ MS provided by Waters Corp. (Milford, Mass.),though other suitable instruments are also commercially available andfrom other suppliers. As is known in the art, both the elution time andthe mass spectrometric response are calibrated using known compoundspresent in known amounts, compounds and amounts that are appropriate tothose occurring in actual unknown samples.

Another embodiment of the current invention relates to partial oxidationof aromatic oxidizable compound with appropriate balancing of thesuppression of noxious aromatic impurities on the one hand against theproduction of carbon dioxide and carbon monoxide, collectively carbonoxides (COx), on the other. These carbon oxides typically exit thereaction vessel in the off-gas, and they correspond to a destructiveloss of solvent and of oxidizable compound, including the ultimatelypreferred oxidized derivatives (e.g., acetic acid, para-xylene, andTPA). The inventors have discovered lower bounds for the production ofcarbon oxides below which it seems the high creation of noxious aromaticimpurities, as described below, and the low overall conversion level areinevitably too poor to be of economic utility. The inventors have alsodiscovered upper bounds of carbon oxides above which the generation ofcarbon oxides continues to increase with little further value providedby reduction in generation of noxious aromatic impurities.

The inventors have discovered that reducing the liquid-phaseconcentrations of aromatic oxidizable compound feed and of aromaticintermediate species within a reaction medium leads to lower generationrates for noxious impurities during the partial oxidation of aromaticoxidizable compound. These noxious impurities include coupled aromaticrings and/or aromatic molecules containing more than the desired numberof carboxylic acid groups (e.g., in the oxidation of para-xylene thenoxious impurities include 2,6-dicarboxyanthraquinone,2,6-dicarboxyfluorenone, trimellitic acid, 2,5,4′-tricarboxybiphenyl,and 2,5,4′-benzophenone). The aromatic intermediate species includearomatic compounds descended from the feed of oxidizable aromaticcompound and still retaining non-aromatic hydrocarbyl groups (e.g., inthe oxidation of para-xylene the aromatic intermediate species comprisepara-tolualdehyde, terephthaldehyde, para-toluic acid, 4-CBA,4-hydroxymethylbenzoic acid, and alpha-bromo-para-toluic acid). Thearomatic oxidizable compound feed and the aromatic intermediate speciesretaining non-aromatic hydrocarbyl groups, when present in the liquidphase of the reaction medium, appear to lead to noxious impurities in amanner similar to that already disclosed herein for dissolved aromaticspecies lacking non-aromatic hydrocarbyl groups (e.g., isophthalicacid).

Set against this need for higher reaction activity to suppress formationof noxious aromatic impurities during partial oxidation of oxidizablearomatic compound, the inventors have discovered that the undesirableattendant result is increased production of carbon oxides. It isimportant to appreciate that these carbon oxides represent a yield lossof oxidizable compound and oxidant, not just solvent. Explicitly, asubstantial and sometimes principal fraction of the carbon oxides comesfrom the oxidizable compound, and its derivatives, rather than fromsolvent; and often the oxidizable compound costs more per carbon unitthan does solvent. Furthermore, it is important to appreciate that thedesired product carboxylic acid (e.g., TPA) is also subject toover-oxidation to carbon oxides when present in the liquid phase of thereaction medium.

It is also important to appreciate that the present invention relates toreactions in the liquid phase of the reaction medium and to reactantconcentrations therein. This is in contrast to some prior inventionsthat relate directly to the creation in precipitated solid form ofaromatic compound retaining non-aromatic hydrocarbyl groups.Specifically, for the partial oxidation of para-xylene to TPA, certainprior inventions pertain to the amount of 4-CBA precipitated in thesolid phase of CTA. However, the present inventors have discovered avariance of greater than two to one for the ratio of 4-CBA in the solidphase to 4-CBA in the liquid phase, using the same specifications oftemperature, pressure, catalysis, solvent composition and space-timereaction rate of para-xylene, depending upon whether the partialoxidation is conducted in a well-mixed autoclave or in a reaction mediumwith oxygen and para-xylene staging according to the present invention.Further, the inventors have observed that the ratio of 4-CBA in thesolid phase to 4-CBA in the liquid phase can also vary by over two toone in either well-mixed or staged reaction medium depending upon thespace-time reaction rate of para-xylene at otherwise similarspecifications of temperature, pressure, catalysis, and solventcomposition. Additionally, 4-CBA in the solid phase CTA does not appearto contribute to the formation of noxious impurities, and 4-CBA in thesolid phase can be recovered and oxidized on to TPA simply and at highyield (e.g., by oxidative digestion of the CTA slurry as is describedherein); whereas the removal of noxious impurities is far more difficultand costly than removal of solid phase 4-CBA, and the production ofcarbon oxides represents a permanent yield loss. Thus, it is importantto distinguish that this aspect of the present invention relates toliquid-phase compositions in the reaction medium.

Whether sourced from solvent or oxidizable compound, the inventors havediscovered that at conversions of commercial utility the production ofcarbon oxides relates strongly to the level of overall reaction activitydespite wide variation in the specific combination of temperature,metals, halogens, temperature, acidity of the reaction medium asmeasured by pH, water concentration employed to obtain the level ofoverall reaction activity. The inventors have found it useful for thepartial oxidation of xylene to evaluate the level of overall reactionactivity using the liquid-phase concentration of toluic acids at themid-height of the reaction medium, the bottom of the reaction medium,and the top of the reaction medium.

Thus, there arises an important simultaneous balancing to minimize thecreation of noxious impurities by increasing reaction activity and yetto minimize the creation of carbon oxides by lowering reaction activity.That is, if the overall production of carbon oxides is suppressed toolow, then excessive levels of noxious impurities are formed, and viceversa.

Furthermore, the inventors have discovered that the solubility and therelative reactivity of the desired carboxylic acid (e.g., TPA) and thepresence of other dissolved aromatic species lacking non-aromatichydrocarbyl groups introduce a very important fulcrum in this balancingof carbon oxides versus noxious impurities. The desired productcarboxylic acid is typically dissolved in the liquid phase of thereaction medium, even when also present in solid form. For example, attemperatures in the preferred ranges, TPA is soluble in a reactionmedium comprising acetic acid and water at levels ranging from about onethousand ppmw to in excess of 1 weight percent, with solubilityincreasing as temperature increases. Notwithstanding that there aredifferences in the reaction rates toward forming various noxiousimpurities from oxidizable aromatic compound feed (e.g., para-xylene),from aromatic reaction intermediates (e.g., para-toluic acid), from thedesired product aromatic carboxylic acid (e.g., TPA), and from aromaticspecies lacking non-aromatic hydrocarbyl groups (e.g., isophthalicacid), the presence and reactivity of the latter two groups establishesa region of diminishing returns with regards to further suppression ofthe former two groups, oxidizable aromatic compound feed and aromaticreaction intermediates. For example, in a partial oxidation ofpara-xylene to TPA, if dissolved TPA amounts to 7,000 ppmw in the liquidphase of the reaction medium at given conditions, dissolved benzoic acidamounts to 8,000 ppmw, dissolved isophthalic acid amounts to 6,000 ppmw,and dissolved phthalic acid amounts to 2,000 ppmw, then the value towardfurther lowering of total noxious compounds begins to diminish asreaction activity is increased to suppress the liquid-phaseconcentration para-toluic acid and 4-CBA below similar levels. That is,the presence and concentration in the liquid phase of the reactionmedium of aromatic species lacking non-aromatic hydrocarbyl groups isvery little altered by increasing reaction activity, and their presenceserves to expand upwards the region of diminishing returns for reducingthe concentration of reaction intermediates in order to suppressformation of noxious impurities.

Thus, one embodiment of the present invention provides preferred rangesof carbon oxides (carbon monoxide and carbon dioxide), bounded on thelower end by low reaction activity and excessive formation of noxiousimpurities and on the upper end by excessive carbon losses, but atlevels lower than previously discovered and disclosed as commerciallyuseful. Accordingly, the formation of carbon oxides is preferablycontrolled as follows. The ratio of moles of total carbon oxidesproduced to moles of oxidizable aromatic compound fed is preferably inthe range of from about 0.02:1 to about 0.25:1, more preferably in therange of from about 0.04:1 to about 0.22:1, still more preferably in therange of from about 0.05:1 to about 0.19:1, and most preferably in therange of from 0.06:1 to 0.15:1. The ratio of moles of carbon dioxideproduced to moles of oxidizable aromatic compound fed is preferably inthe range of from about 0.01:1 to about 0.21:1, more preferably in therange of from about 0.03:1 to about 0.19:1, still more preferably in therange of from about 0.04:1 to about 0.16:1, and most preferably in therange of from 0.05:1 to 0.11:1. The ratio of moles of carbon monoxideproduced to moles of oxidizable aromatic compound fed is preferably inthe range of from about 0.005:1 to about 0.09:1, more preferably in therange of from about 0.01:1 to about 0.07:1, still more preferably in therange of from about 0.015:1 to about 0.05:1, and most preferably in therange of from 0.02:1 to 0.04.

The content of carbon dioxide in dry off-gas from the oxidation reactoris preferably in the range of from about 0.1 to about 1.5 mole percent,more preferably in the range of from about 0.20 to about 1.2 molepercent, still more preferably in the range of from about 0.25 to about0.9 mole percent, and most preferably in the range of from 0.30 to 0.8mole percent. The content of carbon monoxide in dry off-gas from theoxidation reactor is preferably in the range of from about 0.05 to about0.6 mole percent, more preferably in the range of from about 0.10 toabout 0.5 mole percent, still more preferably in the range of from 0.15to about 0.35 mole percent, and most preferably in the range of from0.18 to 0.28 mole percent.

The inventors have discovered that an important factor for reducing theproduction of carbon oxides to these preferred ranges is improving thepurity of the recycled solvent and of the feed of oxidizable compound toreduce the concentration of aromatic compounds lacking non-aromatichydrocarbyl groups according to disclosures of the presentinvention—this simultaneously reduces the formation of carbon oxides andof noxious impurities. Another factor is improving distribution ofpara-xylene and oxidant within the reaction vessel according todisclosures of the present invention. Other factors enabling the abovepreferred levels of carbon oxides are to operate with the gradients inthe reaction medium as disclosed herein for pressure, for temperature,for concentration of oxidizable compound in the liquid phase, and foroxidant in the gas phase. Other factors enabling the above preferredlevels of carbon oxides are to operate within the disclosures hereinpreferred for space-time reaction rate, pressure, temperature, solventcomposition, catalyst composition, and mechanical geometry of thereaction vessel.

One possible benefit of operating within the preferred ranges of carbonoxide formation is that the usage of molecular oxygen can be reduced,though not to stoichiometric values. Notwithstanding the good staging ofoxidant and oxidizable compound according to the present invention, anexcess of oxygen must be retained above the stoichiometric value, ascalculated for feed of oxidizable compound alone, to allow for somelosses to carbon oxides and to provide excess molecular oxygen tocontrol the formation of noxious impurities. Specifically for the casewhere xylene is the feed of oxidizable compound, the feed ratio ofweight of molecular oxygen to weight of xylene is preferably in therange of from about 0.9:1 to about 1.5:1, more preferably in the rangeof from about 0.95:1 to about 1.3:1, and most preferably in the range offrom 1:1 to 1.15:1. Specifically for xylene feed, the time-averagedcontent of molecular oxygen in the dry off-gas from the oxidationreactor is preferably in the range of from about 0.1 to about 6 molepercent, more preferably in the range of from about 1 to about 2 molepercent, and most preferably in the range of from 1.5 to 3 mole percent.

Another possible benefit of operating within the preferred ranges ofcarbon oxide formation is that less aromatic compound is converted tocarbon oxides and other less valuable forms. This benefit is evaluatedusing the sum of the moles of all aromatic compounds exiting thereaction medium divided by the sum of the moles of all aromaticcompounds entering the reaction medium over a continuous period of time,preferably one hour, more preferably one day, and most preferably 30consecutive days. This ratio is hereinafter referred to as the “molarsurvival ratio” for aromatic compounds through the reaction medium andis expressed as a numerical percentage. If all entering aromaticcompounds exit the reaction medium as aromatic compounds, albeit mostlyin oxidized forms of the entering aromatic compounds, then the molarsurvival ratio has its maximum value of 100 percent. If exactly 1 ofevery 100 entering aromatic molecules is converted to carbon oxidesand/or other non-aromatic molecules (e.g., acetic acid) while passingthrough reaction medium, then the molar survival ratio is 99 percent.Specifically for the case where xylene is the principal feed ofoxidizable aromatic compound, the molar survival ratio for aromaticcompounds through the reaction medium is preferably in the range of fromabout 98 to about 99.9 percent, more preferably in the range of fromabout 98.5 to about 99.8 percent, and most preferably in the range offrom 99.0 to 99.7 percent.

Another aspect of the current invention involves the production ofmethyl acetate in a reaction medium comprising acetic acid and one ormore oxidizable aromatic compounds. This methyl acetate is relativelyvolatile compared to water and acetic acid and thus tends to follow theoff-gas unless additional cooling or other unit operations are employedto recover it and/or to destroy it prior to releasing the off-gas backto the environment. The formation of methyl acetate thus represents anoperating cost and also a capital cost. Perhaps the methyl acetate isformed by first combining a methyl radical, perhaps from decompositionof acetic acid, with oxygen to produce methyl hydroperoxide, bysubsequently decomposing to form methanol, and by finally reacting theproduced methanol with remaining acetic acid to form methyl acetate.Whatever the chemical path, the inventors have discovered that whenevermethyl acetate production is at too low a rate, then the production ofcarbon oxides are also too low and the production of noxious aromaticimpurities are too high. If methyl acetate production is at too high arate, then the production of carbon oxides are also unnecessarily highleading to yield losses of solvent, oxidizable compound and oxidant.When employing the preferred embodiments disclosed herein, theproduction ratio of moles of methyl acetate produced to moles ofoxidizable aromatic compound fed is preferably in the range of fromabout 0.005:1 to about 0.09:1, more preferably in the range of fromabout 0.01:1 to about 0.07:1, and most preferably in the range of from0.02:1 to about 0.04:1.

When the generation of carbon dioxide, carbon monoxide, their sum,and/or methyl acetate are below the preferred ranges disclosed herein orwhen the molar survival ratio for aromatic compounds is above thepreferred ranges disclosed herein, the reaction activity should beincreased or the STR should be reduced. One activity accelerator isincreased temperature, within the preferred ranges disclosed herein.Another activity accelerator is increased catalytic activity as providedby the mixture of catalytic chemicals and solvent. Generally, increasingcobalt and/or bromine concentrations will accelerate reaction activity,if these are being used within the ranges preferred herein. Adjustingthe concentration within the reaction medium of other catalystcomponents and of water can also be used to accelerate reactionactivity. STR is decreased by decreasing the feed rate of oxidizablecompound and/or by increasing the volume of reaction medium.

When the generation of carbon dioxide, carbon monoxide, their sum,and/or methyl acetate is greater than the preferred ranges disclosedherein and/or when the molar survival ratio for aromatic compounds isbelow the preferred ranges disclosed herein, preferable control actionsinclude a reverse of the above actions, again within the preferredranges disclosed herein. The inventors note that it is particularlyhelpful to raise the STR as far as possible into the ranges herein whilemaintaining a good quality of oxidation as measured by noxiousimpurities in the CTA and in the reaction medium. The inventors againnote that it is difficult to maintain this quality of oxidation at suchhigh STR and that very careful attention is required with respect towardthe following: to feed dispersion upon entering the reaction medium, toaeration quality throughout the reaction medium, to de-aeration uponexit from the reaction medium, to oxygen-STR and dissolved oxygenthroughout the reaction medium, to excess oxidant exiting the reactionmedium, to the desirable spatial gradient of oxygen-STR, to thedesirable spatial gradient of oxidizable compound concentration, to thedesirable spatial gradient of oxidant concentration, to the overheadpressure, to the desirable spatial gradient of pressure, and to thepreferred temperature at the mid-height of the reaction medium, and asare all disclosed herein. In further addition and in order to achievelower carbon dioxide, carbon monoxide, and/or their sum and/or in orderto increase the molar survival ratio for aromatic compounds, theinventors have discovered that it is useful to suppress within thereaction medium the concentration of soluble aromatic compounds lackingnon-aromatic hydrocarbyl groups (e.g. isophthalic acid, phthalic acidand benzoic acid); this suppression may be effected by using purer feedof oxidizable compound and/or purer solvent, especially within thepreferred ranges for each as disclosed herein.

In a reaction medium continuously oxidizing para-xylene to terephthalicacid at the preferred STR disclosed herein, it is preferred that theamount of para-toluic acid in the liquid phase of the reaction medium bemaintained in the range from about 200 to about 10,000 ppmw, morepreferably from about 800 to about 8,000 ppmw and most preferably from1,600 to 6,000 ppmw. Furthermore, conversion of para-xylene toterephthalic acid within the reaction medium is preferably maintainedabove about 50 mole percent, more preferably above about 90 molepercent, still more preferably above about 95 mole percent, and mostpreferably above 97 mole percent.

In one embodiment of the present invention, it is preferred for one ormore of the operating parameters disclosed herein (includingnumerically-quantified operating parameters) to be maintained for acommercially-significant period of time. Preferably, operation inaccordance with one or more of above-described operating parameters ismaintained for at least about 1 hour, more preferably, at least about 12hours, still more preferably at least about 36 hours, and mostpreferably at least 96 hours. Thus, unless otherwise indicated herein,the operating parameters described herein are intended to apply tosteady-state, optimal/commercial operation—not start-up, shut-down, orsub-optimal operation.

The inventors note that for all numerical ranges provided herein, theupper and lower ends of the ranges can be independent of one another.For example, a numerical range of 10 to 100 means greater than 10 and/orless than 100. Thus, a range of 10 to 100 provides support for a claimlimitation of greater than 10 (without the upper bound), a claimlimitation of less than 100 (without the lower bound), as well as thefull 10 to 100 range (with both upper and lower bounds). Further, whenthe term “about” is used to modify a numerical value, it should beunderstood that in one embodiment, the numerical value is the exactnumerical value.

The invention has been described in detail with particular reference topreferred embodiments thereof, but will be understood that variationsand modification can be effected within the spirit and scope of theinvention.

1. A process for making a polycarboxylic acid composition, said processcomprising: (a) subjecting a multi-phase reaction medium to oxidation ina primary oxidation reactor to thereby produce a first slurry; (b)subjecting at least a portion of said first slurry to further oxidationin a secondary oxidation reactor, wherein said secondary oxidationreactor is a bubble column reactor, wherein said primary oxidationreactor defines therein a primary reaction zone, wherein said secondaryoxidation reactor defines therein a secondary reaction zone, wherein theratio of the volume of said primary reaction zone to the volume of saidsecondary reaction zone is in the range of from about 4:1 to about 50:1.2. The process of claim 1 further comprising introducing an aromaticcompound into said primary oxidation reactor, wherein at least 80 weightpercent of said aromatic compound introduced into said primary oxidationreactor is oxidized in said primary oxidation reactor.
 3. The process ofclaim 2 wherein said aromatic compound is para-xylene.
 4. The process ofclaim 1 wherein step (b) includes oxidizing para-toluic acid present insaid first slurry.
 5. The process of claim 4 further comprisingwithdrawing a second slurry from said secondary oxidation reactor,wherein the time-averaged concentration of para-toluic acid in theliquid phase of said second slurry is less than 50 percent of thetime-averaged concentration of para-toluic acid in the liquid phase ofsaid first slurry.
 6. The process of claim 5 wherein the time-averagedconcentration of para-toluic acid in the liquid phase of said firstslurry is at least 500 ppmw, wherein the time-averaged concentration ofpara-toluic acid in the liquid phase of said second slurry is less than250 ppmw.
 7. The process of claim 1 wherein said primary oxidationreactor is a bubble column reactor.
 8. The process of claim 1 whereinsaid secondary oxidation reactor is located outside of said primaryoxidation reactor.
 9. The process of claim 8 wherein at least a portionof said secondary oxidation reactor is located alongside said primaryoxidation reactor.
 10. The process of claim 1 wherein said secondaryoxidation reactor is not a piston flow reactor.
 11. The process of claim1 further comprising removing said first slurry from said primaryoxidation reactor via a slurry outlet located between the top and bottomends of said primary oxidation reactor.
 12. The process of claim 11wherein said primary oxidation reactor defines therein a primaryreaction zone having a maximum height (H_(p)), wherein said slurryoutlet is spaced at least 0.1 H_(p) from the bottom and top ends of saidprimary reaction zone.
 13. The process of claim 12 wherein said slurryoutlet is spaced at least 0.25 H_(p) from the bottom and top ends ofsaid primary reaction zone.
 14. The process of claim 1 wherein saidprimary reaction zone has a ratio of maximum vertical height to maximumhorizontal diameter in the range of from about 3:1 to about 30:1,wherein said secondary reaction zone has a ratio of maximum verticalheight to maximum horizontal diameter in the range of from about 1:1 toabout 50:1.
 15. The process of claim 1 wherein the ratio of the maximumhorizontal diameter of said secondary reaction zone to the maximumhorizontal diameter of said primary reaction zone is in the range offrom about 0.1:1 to about 0.6:1, wherein the ratio of the maximumvertical height of said secondary reaction zone to the maximum verticalheight of said primary reaction zone is in the range of from about 0.1:1to about 0.9:1.
 16. The process of claim 1 wherein said primary reactionzone has a maximum diameter (D_(r)), wherein the volumetric centroid ofsaid secondary reaction zone is horizontally spaced at least 0.5 D_(p)from the volumetric centroid of said primary reaction zone.
 17. Theprocess of claim 16 wherein said primary reaction zone has a maximumheight (H_(p)), wherein the volumetric centroid of said secondaryreaction zone is vertically spaced less than 0.5 H_(p) from thevolumetric centroid of said primary reaction zone.
 18. A process formaking a polycarboxylic acid composition, said process comprising: (a)subjecting a multi-phase reaction medium to oxidation in a primaryoxidation reactor to thereby produce a first slurry; (b) subjecting atleast a portion of said first slurry to further oxidation in a secondaryoxidation reactor to thereby produce a second slurry comprising crudepolycarboxylic acid particles, wherein said secondary oxidation reactoris a bubble column reactor, wherein the time-averaged concentration ofpara-toluic acid in the liquid phase of said first slurry is at least1,000 ppmw, wherein the time-averaged concentration of para-toluic acidin the liquid phase of said second slurry is less than 1,000 ppmw; and(c) subjecting at least a portion of said second slurry to at least onechemical reaction to thereby produce a third slurry comprising purifiedpolycarboxylic acid particles, wherein said purified polycarboxylic acidparticles comprise a concentration of 4-carboxybenzaldehyde (“4-CBA”)that is at least 200 ppmw less than the 4-CBA concentration of saidcrude polycarboxylic acid particles.
 19. The process of claim 18,wherein said primary oxidation reactor and said secondary oxidationreactor are each operated at a time-averaged and volume-averagedtemperature in the range of from about 140 to about 180° C., wherein theoverhead pressure of said primary oxidation reactor and said secondaryoxidation reactor are each maintained in the range of from about 1 toabout 20 barg.
 20. The process of claim 18, wherein the pressure andtemperature in said secondary oxidation reactor are approximately equalto the pressure and temperature in said primary oxidation reactor. 21.The process of claim 18, wherein said primary oxidation reactor definestherein a primary reaction zone, wherein said secondary oxidationreactor defines therein a secondary reaction zone, wherein the ratio ofthe volume of said primary reaction zone to the volume of said secondaryreaction zone is in the range of from about 4:1 to about 50:1.
 22. Theprocess of claim 21, wherein the ratio of the volume of said primaryreaction zone to the volume of said secondary reaction zone is in therange of from 8:1 to 30:1.
 23. The process of claim 18, wherein saidsecondary oxidation reactor defines therein a secondary reaction zone,wherein the liquid phase of said first slurry has a residence time insaid secondary reaction zone in the range of from about 2 to about 60minutes.
 24. The process of claim 18, wherein said chemical reaction ofstep (c) is performed at a temperature in the range of from about 20 toabout 80° C. higher than the time-averaged and volume-averagedtemperature in said primary oxidation reactor and in said secondaryoxidation reactor.
 25. The process of claim 18, wherein said chemicalreaction of step (c) includes oxidative digestion.
 26. The process ofclaim 18, wherein the total concentration of solids in said first slurryis at least 15 weight percent.
 27. The process of claim 18, wherein saidprimary oxidation reactor is a bubble column reactor.
 28. The process ofclaim 18, wherein said aromatic compound is para-xylene.
 29. The processof claim 1, wherein the liquid phase of said initial slurry has aresidence time in said secondary reaction zone in the range of fromabout 2 to about 60 minutes.
 30. The process of claim 1, wherein saidprimary oxidation reactor and said secondary oxidation reactor are eachoperated at a time-averaged and volume-averaged temperature in the rangeof from about 140 to about 180° C., wherein the overhead pressure ofsaid primary oxidation reactor and said secondary oxidation reactor areeach maintained in the range of from about 1 to about 20 barg.
 31. Theprocess of claim 1, wherein the pressure and temperature in saidsecondary oxidation reactor are approximately equal to the pressure andtemperature in said primary oxidation reactor.